An overview of IGCC systems

An overview of IGCC systems Ting Wang University of New Orleans, New Orleans, LA, United States

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1.1  Introduction of IGCC IGCC is an acronym for Integrated Gasification Combined Cycle. The major purpose of IGCC is to use hydrocarbon fuels in solid or liquid phases to produce electrical power in a cleaner and more efficient way via gasification, compared to directly combusting the fuels. The hydrocarbon fuels typically include coal, biomass, refinery bottom residues (such as petroleum coke, asphalt, visbreaker tar, etc.), and municipal wastes. The approach to achieve a “cleaner” production of power is to convert solid/liquid fuels to gas first, so that they can be cleaned before they are burned by removing mainly particulates, sulfur, mercury, and other trace elements. The cleaned gas, called synthetic or synthesis gas (syngas), which primarily consists of carbon monoxide (CO) and hydrogen (H2), can then be sent to a conventional combined cycle to produce electricity. A simplified IGCC process diagram comprising three major “islands”—gasification, gas cleanup, and power—is shown in Fig. 1.1. The ultimate goal for IGCC is to achieve a lower cost of electricity (COE) than conventional pulverized coal (PC) power plants and/or to be competitive with natural gas-fired combined-cycle systems with comparable emissions. While “clean” power generation is the primary driving motivation for entering the business of IGCC, “increasing plant efficiency” to a level higher than that of PC plants is the second driving motivation. To achieve higher efficiency, “integration” between sub-systems becomes necessary. Integration consists of all aspects of the operation, including mechanical, thermal, and dynamic process control. For example, mechanical integration can be achieved between the compressor of the gas turbine (GT) and the air separation unit (ASU), aiming to save some compression power. Thermal integration can be implemented by strategically interconnecting the various grades of steam generated during the syngas cooling, gas cleanup, and/or watergas shift processes with the heat recovery steam generator (HRSG) and the steam turbine system. Full air integration does enhance the overall plant efficiency positively by about three to four percentage points, but it also increases the complexity of construction, operation, and maintenance, which may result in increased potential for construction phase delay and/or cost overrun, increased maintenance, lost availability, and degraded reliability. Thus, the concept of nonintegrated IGCC has been advocated by some developers to trade reduced efficiency for higher availability and reliability, even though the term “nonintegrated IGCC” could be confusing. When the potential of global warming became a concern, the emission of carbon dioxide (CO2)—a greenhouse gas (GHG)—from power plants was subjected to Integrated Gasification Combined Cycle (IGCC) Technologies. DOI: http://dx.doi.org/10.1016/B978-0-08-100167-7.00001-9 © 2017 Elsevier Ltd. All rights reserved.

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Integrated Gasification Combined Cycle (IGCC) Technologies

Gas Stream Cleanup/Component Separation

Gasifier

Syngas

Syngas

Fuels

CO/H2

Chemicals

Water–Gas Shift CO2

H2

CO2 Pre-combustion Capture

Particulates Fuel Cell

Oxygen

Electrical Power Combined Cycle

Sulfur/Sulfuric Acid Nitrogen Air ASU

Air

Combustion

Turbine

Generator

Electrical Power

Exhaust Cleaned Exhaust

Steam

Marketable Solid By-Products

Heat Recovery Steam Generator

Exhaust

Stack

CO2 Post-combustion Capture

Condenser

Steam Turbine

Generator

Electrical Power

Power Island

Figure 1.1  Simplified block diagram of an IGCC system.

stringent scrutinization and regulations. Usually, there are three ways to reduce CO2 emissions: by increasing the overall system efficiency, capturing a portion of the CO2 and sequestering it, called CCS (Carbon Capture and Sequestration), or utilizing the captured CO2 multiple times. The syngas generated via the gasification process can be more readily separated into highly concentrated H2 and CO2 through the water-gas shift (WGS) process (to be explained later) before the combustion stage (i.e., precombustion) in an IGCC system, as opposed to PC power plants, which have to use a post-combustion carbon capture method. It is significantly cheaper to perform precombustion carbon capture in an IGCC system than post-combustion carbon capture in a PC power plant due to the nature of the processes involved and the reduced size of equipment. CCS imposes a severe penalty on power output, plant efficiency, and COE. The objective of this chapter is to provide an introduction of the complete IGCC system, allowing readers quickly to obtain an overall view of the IGCC system, leaving the details in each subsequent chapters, each focusing on a specific subject. Although the gasification process can be applied to various carbon fuels, since the major developments and applications have involved coal, the descriptions and explanations in this chapter are written with coal in mind as the major feedstock unless specified.

An overview of IGCC systems

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1.2  Layouts of key IGCC components and processes For the convenience of explaining the IGCC systems with information of some of the flow’s thermodynamic properties, the flow system diagrams obtained from an academic simulation of an IGCC plant are used. The simulation was performed using the commercial software, GT Pro, a part of the program suite, Thermoflow. The plant was designed to generate about 240 MWe of net power output, using Texas lignite as feedstock. The results of these simulated IGCC plants have been documented by Wang and Long (2012a, 2012b, and 2014). Two systems were simulated in Wang and Long’s papers. The result of the one with a lower steam turbine inlet pressure (1100 psi/76 bar) and temperature (538°C/1000°F) is used in this chapter. Fig. 1.2 shows the general layout of the baseline case with and without CCS. The feedstock is the South Hallsville Texas Lignite with a feeding rate of 4308 tons/day. The reason of using the Texas Lignite is because the simulate plant is located in Louisiana and Texas Lignite is close by. The coal is mixed with 35% water by weight to form a slurry, which is injected into a GE entrained flow gasifier together with 95% pure oxygen provided by the air separation unit (ASU). The syngas coming out of the gasifier needs to be cooled down to meet the operating conditions of the currently available gas cleanup system. Typically, either a radiant syngas cooler or a quench cooling method can be used, followed by several traditional convective heat exchanger coolers. The gas cleanup system consists of a scrubber to remove particulates and other soluble contaminants, such as hydrogen cyanide (HCN), ammonia (NH3), and hydrochloric acid (HCl). The slight amount of carbonyl sulfide (COS) in the syngas is converted to hydrogen sulfide (H2S) through COS hydrolysis. The syngas needs to be further cooled down to near the ambient temperature before it enters the Acid Gas Removal (AGR) unit. The heat released from the cooling process between the exit of the gasifier and the inlet of the AGR unit is used to generate superheated steam and hot water at various pressures. The cleaned syngas is sent to the GT to generate electricity. The exhaust of the GT is at about 593°C (1100°F), which has sufficient energy to generate steam through a Heat Recovery Steam Generator (HRSG). The steam generated through the HRSG is combined with steam generated through the syngas cooling process to drive a steam turbine and generate additional electricity. This is identical to a conventional combined cycle. In this example here, a GE quench-type gasifier is used. The power block consists of a single GT, modeled after the Siemens SGT6-4000F turbine, with steam injection in the combustor to reduce NOx formation, and a single ST, with a fixed steam inlet pressure and temperature of 1100 psi (76 bar) and 538°C (1000°F), respectively. The steam is reheated to 538°C (1000°F) at 174.5 psi (11.87 bar) to increase the output power and efficiency of the bottom steam cycle. The plant is designed exclusively for power generation, so no chemicals or energy gases are exported anywhere in the middle of cleanup. If carbon capture is needed in a system that was initially designed without considering carbon capture, a post-combustion carbon capture system (shown as an inset in Fig. 1.2) can be implemented at the exhaust gas side exit of the HRSG. The carbon capture system makes use of an amine-based solvent to separate the CO2 from the rest of the GT exhaust. The cost of using a post-combustion carbon capture system is

Figure 1.2  A general layout of a simulated IGCC plant without CCS with an inset showing an added post-combustion carbon capture system. The physical parameters at each nodal point is represented as pressure p (psia), temperature T(oF), enthalpy h(Btu/lb), and mass flow rate M(lb/s) (Wang and Long, 2012a).

An overview of IGCC systems

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typically higher than a corresponding pre-combustion system in terms of dollar/ton CO2, but its operation is expected to be less complicated than a pre-combustion carbon capture system because the operation of post-combustion CCS has a minimal affect on the operation of upstream power-producing devices, whereas the operation of pre-combustion CCS is tightly intertwined with the entire IGCC system. To reduce the cost of carbon capture, IGCC is particularly cost-effective for implementing pre-combustion carbon capture. The major difference between post-combustion and pre-combustion carbon capture lies in the implementation of Water-Gas Shift (WGS) before the syngas is burned in the GT to convert carbon monoxide and steam to carbon dioxide and hydrogen. Then, the carbon dioxide is separated from the hydrogen, so carbon dioxide can be captured effectively and transported for storage. There are two approaches for pre-combustion carbon capture: sour-shift and sweet-shift. As their names indicate, the sour-shift process installs the WGS unit upstream of the AGR unit, before the sulfur is removed from the syngas, so the WGS process occurs in an acidic environment; whereas, the sweet-shift process implements the WGS unit downstream of the AGR unit after the sulfur is removed from the syngas. Fig. 1.3 illustrates a sour-shift process with an inset to show how a sweet-shift scheme can be implemented by replacing the red dashed circle with the blue dash-dotted box.

Detailed Description of Each Process and Component 1.3  Gasification process Gasification is different from combustion. The purpose of combustion is to produce heat, whereas the purpose of gasification is to produce fuels or chemicals. Therefore, during a combustion process, the stoichiometric (or theoretical) amount of oxidant is used to completely oxidize the feedstock and obtain the maximum thermal energy output (heat); whereas, during a gasification process, as little thermal energy as possible is intended to be used (and, thus, limited oxidant is needed) to convert the feedstock to useful fuels, preserving as much of the original fuel’s chemical energy (or heating value) as desired. Typically, a stoichiometric ratio of 0.25–0.35(i.e., 25–35% of the oxygen theoretically needed for complete combustion) is implemented in a gasification process. Since only limited oxidant is needed, the gasification process has been commonly introduced as an incomplete combustion or partial combustion process. Although it is not wrong to say so, it could be misleading because the purpose of incomplete or partial oxidation is to produce heat, which is only the first step. The resulting heat is needed to complete the rest of gasification process. The actual reactions involved with gasification are extremely complicated and vary with the properties of the feedstock. For the convenience of further explaining the gasification process, a set of simplified, major global reactions involved in a gasification process are summarized as follows:

Figure 1.3  An overall layout of a simulated IGCC plant employing sour-shift for carbon capture. For sweet-shift, the red dashed circle is replaced by the blue dash-dotted box and the CO shift is not coprocessed with COS hydrolysis (Wang and Long, 2012b).

An overview of IGCC systems

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Heterogeneous reactions: ∆H°R = -110.5 MJ/kmol

C(s) + ½ O2 → CO

[R1.1]

C(s) + CO2 ↔ 2CO ∆H°R = +172.0 MJ/kmol (Gasification, Reverse Boudouard reaction) [R1.2] C(s) + H2O(g) → CO + H2 C + 2H2 → CH4,

∆H°R = +131.4 MJ/kmol (Steam-Char Gasification)

∆H°R = -87.4 MJ/kmol (Hydrogasificaiton, Direct methanation)

[R1.3] [R1.4]

Homogeneous reactions: CO + ½ O2 → CO2

∆H°R = -283.1 MJ/kmol

CO + H2O(g) ↔ CO2 + H2

∆H°R = -41.0 MJ/kmol

CO + 3H2 ↔ CH4 + H2O

∆H°R = -205.7 MJ/kmol (Methanation)

[R1.5] (Water-gas shift)

CHmOnNoSpClq→aCO+bH2+cCH4+dC2H2+eN2+fHCl+gH2S+hCOS (Volatile cracking)

[R1.6] [R1.7] [R1.8]

CH4+ ½ O2 → CO + 2 H2

∆H°R = -35.7MJ/kmol (Volatiles gasification via CH4)

C2H2 + O2 → 2 CO + H2

∆H°R = -447.83 MJ/kmol (Volatiles gasification via C2H2) [R1.10]

H2 + ½ O2→ H2O

∆H°R = -242MJ/kmol

[R1.9]

[R1.11]

where (a) all the reaction heats, ΔH°R, are based on 298K and 1 atm; (b) “+” indicates endothermic (absorbing heat), and “−” indicates exothermic (releasing heat); (c) heterogeneous reactions are reactions between different phases (here, it represents coal particles reacting with various gases); (d) homogeneous reactions are reactions occurring entirely within the gas phase; and (e) reactions [R1.8], [R1.9], and [R1.10] consist of a simplified two-step thermal cracking gasification model proposed by the author. The gasification of coal particles involves three major steps, as shown in Fig. 1.4: (a) thermal decomposition/pyrolysis (demoisturization and devolatilization), (b) thermal cracking of the volatiles, and (c) char gasification.

1.3.1 Pyrolysis Coal particles undergo pyrolysis when they enter the hot combustion environment, which needs to be created by burning other gaseous or liquid fuels during the ignition process in the beginning. The hot environment needs to be sustained continuously by the heat released from exothermic processes, mainly via [R1.1] and [R1.5]. When the particle temperature reaches the water boiling point, moisture within the coal (i.e., the inherent moisture) vaporizes through a demoisturization process, and leaves the coal’s core structure by migrating to the core surface as steam. The volatiles are then released as the particle temperature continues to increase. This volatile-releasing process is called devolatilization. The word pyrolysis was derived from the Greek: pyro means “fire” and lysis means “separating.” Theoretically, pyrolysis is defined as a thermo-chemical decomposition process of organic material at elevated temperatures in the absence of oxygen. In the

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Integrated Gasification Combined Cycle (IGCC) Technologies

Figure 1.4  Simplified global gasification of coal particles (sulfur and minerals are not included).

gasification process, although oxygen is present, only limited oxygen is supplied and the oxygen is mostly depleted via carbon combustion ([R1.1, 1.5, 1.9, and 1.10]). Thus, the kinetics and phenomena occurring in the pyrolysis process during the gasification process are close to those observed in an environment without oxygen. The mechanics of the pyrolysis process are affected by the physical properties of the char. During the fast heating of the coal particles, the heat transfer coefficient often decreases. This reduced heat transfer rate to the particle surface results in a temperature plateau of about 400°C (752°F) and lasts throughout the devolatilization process. When fast pyrolysis occurs concurrently surrounded by a combustion flame, this is referred to as fast flaming pyrolysis.

1.3.2 Devolatilization Devolatilization is a decomposition process that occurs when, under heating, volatiles are driven out from a hydrocarbon material (like coal). The rate of devolatilization is influenced by temperature, pressure, residence time, particle size, and coal type. The heating causes chemical bonds to rupture, and both the organic and inorganic compounds to decompose. In a typical fixed bed reactor, the process starts at a temperature of around 100°C (212°F) with desorption of gases, such as water vapor, CO2, CH4, and N2, which are stored in the coal pores. When the temperature reaches above 300°C (572°F), the released liquid hydrocarbon called tar becomes important. Gaseous compounds, such as CO, CO2, and steam are also released. When the temperature is above 500°C (932°F), the fuel particles are in a plastic state where they undergo drastic changes in size and shape. The coal particles then harden again and become char when the temperature reaches around 550°C (1022°F). In general, the larger the particle size, the smaller the volatiles yield because, in larger particles, more volatiles may crack, condense, or polymerize, with some

An overview of IGCC systems

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carbon deposition occurring, during their migration from inside the particle to the particle surface. At high pressures, volatiles yields of bituminous coals decrease due to the low vapor pressure of tar. In contrast, low rank coals do not show decreased volatile yields with increased pressure since these coals do not have as much tar.

1.3.3  Volatiles cracking Volatile matter usually consists of a mixture of short- and long-chain hydrocarbons, aromatic hydrocarbons, and some sulfur and chlorine. Volatile matter with longer hydrocarbon chains and higher boiling temperature usually becomes tar, which condenses easily and can cause severe operating problems by plugging the piping and fouling the surfaces of downstream components if the gas temperature decreases below the condensation point. Considering low temperatures (38–138°C/100–280°F) are required during the desulfurization and mercury removal processes downstream, long-chain volatiles need to be cracked into lighter gases. Cracking can usually be performed either catalytically at lower temperatures or thermally at higher temperatures. For IGCC applications, usually thermal cracking is employed to break heavier long-chain volatiles and hydrocarbons (such as C3H8, C6H5OH, C6H6, or substances with longer C-chains) into lighter gases such as H2, CO, C2H2, CH4, C2H6, etc. These lighter gases may react with limited O2 via partial oxidation [R1.9] and [R1.10], releasing more heat, which is needed to continue the pyrolysis, devolatilization, and thermal cracking processes. Reaction [R1.8] models thermal cracking of volatiles in a general form (CHmOnNoSpClq) into CO, H2, CH4, C2H2, N2, HCl, H2S, and COS, followed by a partial oxidation process to convert the intermediate gases, CH4 and C2H2 (the two lightest hydrocarbons), into the desired components, CO and H2, via [R1.8] and [R1.9]. This two-step model removes the need to accurately know the real reaction mechanism (which is usually more complex than the simplified two-step model) during the thermal cracking process because eventually the modeled intermediate gases would be converted to CO and H2 as long as there is O2 available. The temperature in the gasification process is sufficiently high so that, if any hydrocarbons can survive, it would be only the lightest hydrocarbons, for example CH4, that can be present in a noticeable amount. Furthermore, this two-step model also provides a pathway for volatiles to generate heat for the gasification process via incomplete combustion (or partial oxidation). For example, thermal cracking [R1.8] for one category of Illinois No.6 bituminous coal and West Kentucky No.11 lignite can be modeled as (Wang et. al., 2014):

CH 2.761O0.264 N 0.055S0.048 Cl0.005 → 0.256CO + 0.466H 2 + 0.33CH 4 + 0.2C2 H 2 + 0.0275N 2 + 0.005HCl + 0.04H 2 S + 0.008COS (Illinois No.6)

CH3.187 O0.336 N 0.06 S0.01 → 0.8575H 2 + 0.334CO + 0.264CH 4 + 0.2C2 H 2 +00.03N 2 + 0.008H 2 S + 0.002COS (West Kentucky No.11)

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Integrated Gasification Combined Cycle (IGCC) Technologies

1.3.4  Endothermic steam-char and carbon dioxide-char gasification processes With only char and ash left, the char particles undergo two important endothermic, heterogeneous gasification reactions: one is the Boudouard reaction: C(s) + CO2 ↔ 2CO [R1.2] (or, more accurately, the reverse Boudouard reaction, and it is also called as carbon dioxide-char gasification), and the other one is C(s) + H2O(g) → CO + H2 [R1.3] (also called steam-char gasification). Both are endothermic. The Boudouard reaction was discovered in 1905 by the French chemist, Octave Leopold Boudouard (1872–1923), who investigated the equilibrium behavior of the Boudouard reaction, C(s) + CO2 ↔ 2CO (Holleman et al., 2001). When temperature becomes higher than 700°C (1292°F), the reaction is endothermic and tends toward production of CO. Inside a typical gasifier of an IGCC system, in a reducing environment with temperatures higher than 900°C (1652°F), the production of CO becomes dominant (Hunt et al., 2013). The reactivity of the char during the Boudouard reaction can be affected by the catalytic effect of inherent minerals contained in the coal. Typically, the greater the alkali index, the higher the reactivity (Zhang et al., 2003). The Boudouard reaction performs a very heroic action during the gasification process because it converts the villain CO2 into CO, which is a good fuel, as well as a useful chemical feedstock for producing other chemicals, such as methanol or substitute natural gas (SNG). The Boudouard reaction can also be used for CO2 remediation or utilization, but it must be implemented in a high-temperature environment, making the process expensive. The steam-char reaction is the major contributor to the production of both H2 and CO, which are the primary components of the syngas. To take advantage of the steamchar reaction in dry-feed applications, injection of an adequate amount of steam at the appropriate location in the gasifier becomes an important design consideration. The steam-char reaction is about an order of magnitude faster than CO2-char (the Boudouard) reaction. When these two endothermic gasification reactions complete, the major components of the syngas, H2 and CO, have been harnessed. The next important process is the water-gas shift process.

1.3.5  Water-gas shift (WGS) process inside gasifiers The water-gas shift reaction is an equilibrium process: CO + H2O(g) ↔ CO2 + H2 [R1.6]. The forward reaction is exothermic (ΔH°R = −41.0 MJ/kmol), converting carbon monoxide and steam to hydrogen and carbon dioxide. The forward reaction is active at temperatures lower than 700°C. At higher temperatures, near 1000°C, the net reaction is slow and negligible. Beyond 1200°C, the backward reaction becomes dominant. The reaction rate of the WGS is typically slow without using catalysts; however, in the gasifier, the reaction rate is usually enhanced by the catalytic effect of metal components in the coal. Since the forward reaction occurs at relatively low temperatures, the WGS typically occurs in the region in the gasifier where the temperature is reduced due to the endothermic steam-char and carbon dioxide-char gasification processes. The WGS is an important process that will affect the final composition

An overview of IGCC systems

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of the raw syngas. Thus, for dry-feed gasifier operation, a coarse manipulation of the CO/H2 ratio can be achieved by managing the amount of steam being injected into the reduction region of the gasifier. In the gasifiers that use the quench method to cool the syngas down to near 200°C, the residence time is too short to achieve any pronounced forward WGS reaction, even though the equilibrium constant value is high at low temperatures, because the catalytic effect from metals in the coal is minimal in the quenched syngas since most of the metals have become molten slag, which is extracted during the gasification process itself, before quench occurs. For the water-gas shift reaction, the equilibrium constant can be defined as follows: K WGS =



[ H2 ][ CO2 ] (1.1) [ H2 O ][ CO ]

where the brackets represent the concentrations of each chemical compound. For an ideal gas, KWGS has no pressure dependence, and KWGS can be found using the Van’t Hoff equation. ln K WGS = −



∆H ∆S + (1.2) RT R

Here, R is the ideal gas constant, ΔH is the reaction enthalpy, and ΔS is the change in entropy of the reaction. Using the software program Gaseq, the value of KWGS ranging from room temperature to over 1204°C (2200°F) is calculated and shown in Fig. 1.5. A good approximation using the Arrhenius form can be derived as (Fig. 1.6):

 5361 K WGS = 2.51 × 10−6 * T 1.14 * exp  S (T in K ) (1.3a)  T 



 2978.33  K eq,WGS = 4.906 × 10−5 *T 1.14 * exp  , (T in ºR) (1.3b)  T  or by fitting the data in logarithmic form log K WGS = 3627.7(1/T ) − 1.862, (T in ° R) (1.3c)



Some select values at different temperatures are shown in Table 1.1. (B)

1,00,000 90,000 80,000 70,000 60,000 50,000 40,000 30,000 20,000 10,000 0

120 100 Eq. constant, Keq

KWGS

(A)

80 60 40 20

0

500

1000

1500

2000

KWGS Temperature, T (°F)

2500

0 460

960

1460

1960

Temperature, T (°F)

Figure 1.5  WGS equilibrium constant value (A) between 27°C and 1127°C (80°F and 2060°F) (B) zoom in between 238°C and 1127°C (460°F and 2060°F).

2460

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Integrated Gasification Combined Cycle (IGCC) Technologies

6

Eq. 1.3b

log10 (KWGS)

5

Liner Curve Fit Eq. 1.3c

4

log (KWGS) = (3627.7/T) - 1.862

3 2 1 0 –1

0

0.0005

0.001

1/T,

0.0015

0.002

°R–1

Figure 1.6  WGS equilibrium constant value vs. 1/T in semi-log coordinates. Table 1.1  Water-gas T(K) T(°F) log(KWGS)

shift reaction equilibrium constant (KWGS)

400 260 3.194

600 620 1.447

800 980 0.620

1000 1340 0.148

1200 1700

1400 2060

−0.150

−0.351

Lu and Wang (2013a, b) conducted a numerical simulation to investigate the effect of WGS reaction rate on the resultant syngas composition. They compared the simulated result with the experimental results and stated that the available published WGS reaction rates in the open literature were too fast because most of them were obtained from the laboratory condition with different catalytic effects rather than those actually experienced in real gasifers. The WGS is an important process to manipulate the CO/H2 ratio for downstream applications, such as for producing SNG or other chemicals, or for carbon capture applications. More details of downstream WGS applications will be introduced in Section 1.4.

1.3.6 Methanation The two methanation reactions [R1.4 and R1.7] are highly exothermic and pressurefavorable, so they are generally not active in high-temperature environments, such as in high-temperature entrained flow gasifiers. If the goal of gasification is to produce Substitute Natural Gas (SNG), it is usually produced downstream in a low-temperature, high-pressure reactor facilitated with catalysts, mainly consisting of ruthenium, cobalt, nickel, and/or ironbased metals. It would be better if methanation could be directly performed in the gasifier, but current technology has not yet reached that status. Lu and Wang (2016) conducted a review of coal-to-SNG methods and predicted that it would be difficult to produce SNG directly with a methane concentration above 18% (vol) using a once-through gasifier.

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1.4 Gasifiers With the understanding of the fundamental, global gasification process, many different gasifiers have been designed to best achieve certain target syngas compositions (predicted by the equilibrium analysis) with the goals of minimizing the gasifier’s size (thus reducing cost), maximizing output yield, enhancing gasification efficiency, lowering maintenance frequency, and increasing reliability and availability. In order to obtain the syngas composition predicted by chemical equilibrium theory, the residence time must be greater than that calculated from the reaction kinetics, which is typically implemented by making the gasifier large enough or reducing the flow rate, as seen in the design of a fixed bed gasifier. In a fixed bed gasifier, the feedstock is typically fed from the top and falls downwards by gravity. During the free-fall period, pyrolysis and devolatilization occur; any unfinished process is completed in the particle bed in the bottom of the gasifier. There is plenty of time (i.e., long residence time) for the reactions to complete in the sitting particle bed. To increase the contact surface area, the bed is moving either linearly or in a rotating motion with turbulators, or the bed is fixed and is disturbed by a stirrer. Lurgi and British Gas/Lurgi gasifiers (Fig. 1.7) are typical representatives of moving bed gasifiers. This is why fixed bed gasifiers are often also referred to as moving bed gasifiers if some agitation or stirring actions on the bed are added. Thanks to the long residence time, the demand for oxidant is low compared to other types of gasifiers. The coal feed is typically ground down to a size of about 50 mm. The difference between the Lurgi and British Gas/Lurgi (BGL) gasifiers is that the Lurgi gasifier operates in a non-slagging, dry-bottom mode, whereas the BGL gasifier operates in slagging mode. The operating temperatures of updrafting moving bed gasifiers range from 1000°C (1832°F) at the bottom to 540°C (1004°F) at the top, with pressures ranging from 20 to 30 bar (290–435 psig). Fixed bed gasifiers are simple to build. Since the stirrer helps to increase the particle contact surface area, more agitation of the particle bed will be more desirable and appealing. This leads to the design of bubbling fluidized bed gasifiers—basically, the particle bed is agitated into a fluid-like state by injecting air or gas from the bottom of the bed. The air/gas jets travel upward against the weight of the bed particles in pockets, like bubbles in boiling water—only the body of water is replaced with packed particles. Thus, the name bubbling fluidized bed is derived. Since the agitating motion provides more effective contact surface area for heat/mass transfer and chemical reactions for the same syngas yield, the sizes of fluidized bed gasifiers are typically smaller than those of moving bed gasifiers; or, for the same size, fluidized bed gasifiers typically have higher yield. The GTI (U-Gas) gasifier is an example of a bubbling fluidized bed gasifier (Fig. 1.7). To further increase the throughput of the gasifier, the superficial velocity of the gas is increased. The superficial flow velocity is an averaged flow velocity calculated by assuming that the interested gas or fluid phase is the only one flowing or present in a given cross-sectional area, and other phases and particles are absent. The result of the increased gas speed is a significantly increased attrition rate of unreacted chars through the free-board region of the gasifier. Thus, recapture and recycling of the unreacted chars becomes an important mechanism to achieve high carbon conversion.

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Integrated Gasification Combined Cycle (IGCC) Technologies

A cyclone is typically installed to recapture relatively large-sized particles, such as unburned char, soot, and ash. The addition of a faster recycling path yields the name circulating fluidized bed gasifier. The High-Temperature Winkler (HTW) gasifier is an example of a circulating fluidized bed gasifier. The coal feed is typically ground down to a size of about 6 mm (1/4 inches). Fluidized bed gasifiers typically operate at pressures ranging from 35 to 60 bar (508–870 psig) and temperatures ranging from 800°C (1472°F) to1000°C (1832°F). Since the operating temperature range in fluidized bed gasifiers is lower than the ash melting temperatures, these gasifiers produce ash rather than slag. The KBR transport gasifier (Fig. 1.7) used in the Kemper IGCC plant can be also categorized as a circulating fluidized bed. One of the advantages of the KBR

Figure 1.7  Various commercial gasifiers (Stiegel, 2009).

Figure 1.7  (Continued)

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Integrated Gasification Combined Cycle (IGCC) Technologies

transport gasifier is that the return of the recycled char particles is not accomplished using traditional fluid flow. Rather, they are transported instantly by using a gravity wave through accumulated unreacted chars stored in a stand-pipe. Although a char particle isn’t actually transported back to the fluidized bed by itself, when it drops from the cyclone to the top of the stand-pipe, another char particle at the bottom of the stand-pipe can be almost instantaneously pushed out to the fluidized bed at an effective speed equal to that of the gravity (or pressure) wave. Furthermore, the mass flow rate of char recycling can be easily controlled by using the stand-pipe as a carbon carryover “capacitor.” If gas speed is further increased to a point where the fluidized bed can’t be sustained, all of the coal particles will be carried away by the gas—this is how an entrained flow gasifier works. The primary purpose of an entrained flow gasifier is to increase both the throughput and yield significantly. The gas speed can reach 80–100 m/s, and the residence time is typically between 3–5 seconds. Due to the short residence time, more oxidizing agent is needed in a typical entrained flow gasifier than in fluidized beds. The coal feed is usually ground into powder with a particle size of around 0.15 mm. Entrained flow gasifiers operate at pressures ranging from 20 to 70 bar (290–1015 psig) and temperatures ranging from 1250°C (2282°F) to 1600°C (2912°F). Since this operating temperature range is higher than the ash-melting temperatures, entrained flow gasifiers produce slag rather than ash. Theoretically, an entrained flow gasifier can be made large enough so that the reactions will have sufficient residence time to achieve the results predictable through chemical equilibrium calculations. However, due to the non-uniformity of local thermal-flow behavior (or patterns) inside the gasifier, the product may not reach the equilibrium condition at the gasifier’s exit. For example, if there is a large recirculation zone in the gasifier blocking part of the through-flow passage, part of the flow will be re-circulated several times before moving to the gasfier’s exit. This part of the flow will travel for longer than the required residence time; whereas the other part of the flow could be accelerated through the narrowed passage and exit the gasifier in a time period shorter than the required residence time. Therefore, understanding of the thermal-flow field becomes essentially important, particularly when the residence time is shortened because the size of the gasifier is reduced for the purpose of reducing the manufacturing cost of the gasifier. Parekh (1982) published a handbook, providing a comprehensive review of many industrial gasifiers and gas-treatment systems. Basu (2006) also provided a wealth of information of gasification and combustion process in fluidized beds.

1.4.1  Various operation methods of gasifiers As mentioned earlier, the gasification process aims to produce fuels or chemicals by utilizing the minimum amount of thermal energy (heat). Hence, one of the important design considerations is to determine how the heat is created via combustion without raising the gas temperature too high, and how the heat can be distributed strategically to achieve the optimum fuel conversion. This goal leads to the different designs of coal feeding methods: whether the feedstock travels against or in parallel with the gas

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flow direction, whether the feedstock is tangentially injected or through opposing jets, and whether the combustion process is concentrated in a well-protected combustor or strategically distributed in multiple stages. Each of these different operation methods are discussed as follows.

1.4.1.1  By feeding conditions (slurry vs dry) The water-coal ratio is typically between 40 and 60% by weight. Slurry-fed gasifiers are typically more tolerant of variation of coal quality and rank. The IGCC cycle efficiency of slurry-fed gasifiers usually suffers slightly because high-grade energy produced from combustion is used to evaporate the water. On the other hand, for dryfed gasifiers, low-grade heat (such as waste heat from the GT or from the water-gas shift process) can be used to generate steam, which can be injected into the gasifier to enhance the steam-char gasification process [R1.3]. This would increase the cycle efficiency.

1.4.1.2  By oxidizing agent (air-blown vs oxygen-blown) The decision to use oxygen or air as the oxidizing agent is based on three primary considerations: the required heating value of the syngas, the capital and maintenance costs, and power consumption of the ASU. The Higher Heating Value (HHV) of an oxygen-blown syngas ranges from 9,000 to 13,000 kJ/Nm3 (242–350 Btu/scf), which is about 30–35% of natural gas’s HHV (39,000 kJ/Nm3 or 1,047 Btu/scf) and about 2.5–3 times more than that of the syngas produced from air-blown gasifiers. Thus, the mass flow rate of oxygen-blown syngas, without nitrogen in the gas stream, requires smaller gas cleanup and CO2 separation units. However, the capital cost of the ASU island is about 10% of the total cost of an IGCC plant (with no carbon capture) and the auxiliary power consumption is about 13% of the gross power produced by the IGCC plant. (Long and Wang, 2011) These costs and power consumption can be saved if the air-blown gasification process is adopted, resulting in an increased thermal efficiency of approximately 2–3 percentage points. However, the air-blown method will cost more to capture CO2 downstream if CCS is implemented. The Nakoso IGCC plant in Japan and the Kemper IGCC plant in Mississippi are two commercial IGCC plants operating in air-blown mode. Nonetheless, if the syngas is to be used for producing chemicals, the purity and composition of the syngas is important, so the oxygen-blown process is preferred over the air-blown process.

1.4.1.3  By flow directions (counter-flow vs parallel-flow): Counter-flow gasifier—In a counter-flow gasifier (Fig. 1.8), the feedstock is fed from the top and the oxygen/air is fed from the bottom. The cold feedstock from the top is met by hot gases produced further down, which helps dry the feedstock further while cooling down the uprising synthesis gas produced (H2 & CO). This counter-flow heat exchange through cooling of hot syngas by preheating the cold feedstock demonstrates a good use of heat recovery. Next, the very dry falling feedstock is pyrolyzed by the up-flow of hot gases, giving up the volatile combustible gas/vapors and becoming char.

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Integrated Gasification Combined Cycle (IGCC) Technologies

These hot chars are a good reducing agent, as they come into contact with the rising hot CO2 and H2O to yield CO and H2 via gasification reactions [R1.2] and [R1.3]. In the last stage, the char is combusted with the oxygen/air at very high temperatures. The resulting ash is collected and disposed. The advantages of this counter-flow gasifier is the ability to produce slag more readily than the parallel-flow gasifer because metals and minerals are melted due to the high-temperature at the bottom of the gasifier, and the ability to handle very wet fuels. On the other hand, the syngas may contain tars, especially if the feedstock is biomass. The tar is detrimental to GTs and boilers; thus, the tar must be removed before being fed into the GT. In order to avoid the production of tars, the operating temperature in the upper region of the gasifier needs to be raised so that the tars can be thermally cracked into lighter gases. This can be achieved by increasing the operating pressure, reducing the feedstock feeding rate, or increasing the air or oxygen flow rate. The counter-flow gasifier requires good temperature control and monitoring to avoid tar formation and melting of the grate at the bottom of the gasifier, if one is present. The fixed bed gasifiers, Lurgi, and British Gas/Lurgi gasifiers, belong to this category. Parallel-flow gasifier—In parallel-flow gasifiers (Fig. 1.9), both the feedstock and oxidizing agent flow in the same direction; either they all flow from top to bottom or from bottom to top. All the major entrained flow coal gasifiers are operated in parallel-flow mode; they either operate using a downdraft (such as the GE, Simens, or ECUST gasifier) or updraft scheme (such as the E-Gas, Shell, Prenflow, or MHI gasifier.) Since the devolatilization and fast pyrolysis processes occur during the hot combustion process, almost all of the tar is thermally cracked or burned. Hence, the parallel-flow gasifier is more desirable than the counter-flow gasifier when feeding low rank coals or biomass with high-volatile-matter content. Fig. 1.9 illustrates the temperature and producer gas composition changes in a parallel-flow, downdraft gasifier. Near the upper part of the gasifier, the combustion process, fueled by the volatiles, basically consumes most of the oxygen and produces CO2 and steam. As the oxygen is being consumed, the high-temperature combustion gas induces fast flaming pyrolysis, resulting in the production of char. With the absence of oxygen, this leaves char effectively to be gasified with CO2 and steam to produce CO and H2. Since no predrying process occurs in the parallel-flow gasifier, the feedstock intake must have less than 20 wt.% of inherent moisture content. In Shell’s gasifier, a moisture content of the feedstock lower than 12 wt.% is recommended. (US DOE, 2011). Although the feedstock shown in Figs. 1.8 and 1.9 is biomass, the fundamental physics is similar to that of coal gasifiers.

1.4.1.4  By ash melting condition (slagging vs non-slagging) Disposal of coal ash has been always an important issue to keep the environment clean. A non-slagging gasifier disposes of the ash at the bottom of the gasifier through a moving grate in fixed bed gasifiers and controls this process through the use of an ash lock. A slagging gasifier operates with the combustion zone reaching a temperature higher than the ash melting temperature. The molten ash forms a layer on the gasifier’s inner surface and flows down to a water bath at the bottom of the gasifier.

Figure 1.8  Illustration of temperature syngas production process in a counter-flow gasifier (Reed and Gaur, 2001).

Figure 1.9  Example of the temperature and species distribution in an (A) counter-flow gasifier and (B) parallel gasifier fed by biomass (Reed and Gaaur, 2001).

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The molten ash is quenched in the water bath and is solidified as a glass-like, nonleachable solid with all of the metals encapsulated inside. The slag can be sold and used as a safe material for road construction, subject to the local waste disposal regulations. Thus, a non-slagging gasifier is usually dry at the bottom, while a slagging gasifier is wet at the bottom. In non-slagging gasifiers, coals with low ash fusion temperatures may encounter local hot spots and partially melt, becoming sticky and clinging together as sintered clinker particles (Matjie et al., 2008). Agglomerated ash clinkers can cause channel burning, pressure drop problems, and unstable gasifier operation (Van Dyk et al., 2009). Therefore, for non-slagging moving bed and fluidized bed gasifiers, the operating temperature should be much lower than the ash fusion temperature with a sufficient safety margin. A typical issue related to slagging gasifiers is the flowing capability of the molten slag. For gasifiers having a slag tap to guide the slag flowing and dropping to the slag quench bath, the common practice has been using a feedstock with a slag tapping viscosity of 5 Pa·s for slagging fixed bed gasifiers and 15 Pa·s for entrained flow gasifiers. The upper limit for the slag tapping viscosity is 25 Pa·s at a tapping temperature of 1773K (Patterson and Hurst, 2000). If the viscosity of the molten slag is too high, the molten slag layer would flow down slowly, and, eventually, part of the layer adjacent to the wall would solidify, resulting in a thick slag layer, which would literally render the gasifier useless due to the loss of flow space in the gasifier. To overcome this problem, flux, such as limestone, calcites, dolomites, calcium oxide (CaO) or iron(II) oxide (FeO) can be added to reduce the viscosity (or increase the fluidity) of the molten slag (Wang and Massoudi, 2013). These flux materials often can also reduce the ash fusion temperature. Another practice for avoiding slag tap blockage is to adjust the gasifier operating temperature to above the so-called T250 temperature, which is the temperature at which the slag viscosity is 250 centipoise (Barnes, 2013). Another operation issue in slagging gasifiers is related to the plugging problem in the slag tap at the bottom of the gasifier, particularly for updraft entrained flow slagging gasifiers. The slag tap is shaped like a tapered throat; typically the diameter of the opening (throat) is designed about 1/3rd of the inner diameter of the gasifier. The major function of the tap is to increase gasification performance by reducing the heat losses through radiation and convective heat transfer between the gasification chamber and the water bath at the bottom of the gasifier. However, since the wall temperature near the tap is low, and the tap opening is only 1/9th of the cross-sectional area of the gasifier, the molten slag layer may freeze up and clog the throat. To mend this problem, some gasifiers install a torch gun below the tap, which can be ignited to melt the solidified slag in the throat when the throat gets clogged. Another option is to open up the throat opening area at the expense of increased heat losses. Wang et al. (2007) conducted a simulation to investigate the trade-off between the heat losses and the reduced occurrence of clogging the tap by enlarging the tap opening area in a dry-fed, updraft entrained flow gasifier. They discovered that the heat losses could increase by 6% if the tap opening is increased from 1/3rd to 2/3rds of the gasifier inner diameter. If the tap is completely removed, the heat losses increase to 18.3%. The syngas’s LHV values suffer from a reduction of about 3% and 9%, respectively. Changing the

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fuel injection angle from horizontal to downward 15 degrees towards the slag tap can reduce the possibility of clogging with only a slight penalty in increased heat losses and reduced syngas LHV value. Fig. 1.10 shows the flow pathlines and temperature distribution of different slag tap opening areas. Downdraft, slagging gasifiers don’t have this problem because the hot syngas also flows through the slag tap, keeping the molten slag layer from solidifying (Table 1.2). An interesting case of transitioning from a non-slagging design to a slagging design took place when British Gas converted the original non-slagging Lurgi gasifier into a slagging gasifier (Brooks et al., 1984). The slagging operation resulted in a reduction of steam and oxygen consumption and enhanced the quality of the syngas, with more CO and H2 but less unreacted volatiles such as tars, ammonia, and other heavier hydrocarbons (Lohmann and Langhoff, 1982).

1.4.1.5  By wall structure (insulation vs cooling) Three different wall structures have been designed for commercial gasifiers: refractory bricks, water jackets, and membrane walls.

Figure 1.10  Simulated flow pathline traces and temperature distributions of an updraft entrained flow coal gasifier for four cases: (A) d/D = 1/3, (B) d/D = 2/3, (C) d/D = 2/3, 15-degree downward fuel injection, and (D) no slag tap. The fuel is tangentially injected through three burners in a horizontal plane except in Case C. D is the gasifier internal diameter and d is the tap throat diameter. The background color is the translucent back wall temperature, not the gas temperature. The gas temperature is represented by the colors of the dots. Minimal flow pathline traces are seen in the quench section (Wang et al., 2007). Table 1.2 

Effect of slag tap throat size on heat losses and gasification performance (in terms of syngas LHV) Cases

A

B

C

D

Heat losses compared with Case A Syngas LHV compared with Case A

0 0

+6.0% −3.8%

+6.4% −2.5%

+18.3% −8.9%

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Integrated Gasification Combined Cycle (IGCC) Technologies

Due to the high-temperature oxidation in the gasifier, the gasifier wall needs to be insulated to reduce heat losses as well as to protect the gasifier vessel shell and corresponding supporting structure. The refractory used for the gasifier depends on the area of reaction. Near the fuel injector area, high-temperature combustion occurs, so high-temperature resistant refractory bricks made of fusion-cast poly-crystalline materials are typically used. In the area of reduction where the endothermic gasification reactions occur, the temperature is relatively low, and refractory bricks made of chromium oxide- or zirconium oxide-based materials are used. The E-Gas and GE gasifiers are typical examples of reactors that use refractory brick walls. While using refractory bricks can minimize heat losses, corrosion, and erosion on the refractory due to particle-laden flow, flowing slag in the high-temperature and high-pressure harsh environment exerts a toll on the life expectancy of the refractory bricks. The period between maintenance calls for repairing the refractory bricks in the combustion zone has been reported to be between 18 and 24 months. To avoid shutting down the gasifier for refractory maintenance, different methods for cooling the wall have been developed. The Siemens gasifier uses a water jacket to cool the refractory lining for low or zero ash feeds without slagging. The water jacket containing saturated boiling water and steam can be easily and economically constructed as long as the pressure in the water jacket is in equilibrium with that inside the gasifier so no additional structure enforcement is needed. This pressure equilibrium is typically done through communication with the steam flow pathway downstream of the gasifier when the steam in the water jacket is superheated. One advantage of the water jacket is its simplicity and low manufacturing cost; however, the steam temperature is low because it is limited by the pressure of the gasifier. At a pressure 35 bar (508 psia) in a typical commercial gasifier, the steam that can be generated in the water jacket is limited to the saturation temperature of 242.6°C (468.7°F), which is low-grade energy. This means that cooling of the gasifier wall creates a lot of entropy, and the Second Law of Thermodynamics tells us that the heat recovery efficiency of the water jacket is low, and the overall IGCC plant thermal efficiency will be penalized for this inefficient heat recovery. This issue can be resolved by using a cooling screen (or membrane wall), which replaces the water jacket with high-pressure steam tubes. For high-ash feeds, the Siemens gasifier further installs a spirally-wound cooling screen (Fig. 1.11) to replace the refractory lining. The surface of the cooling screen is covered by silicon carbide (SiC) castable refractory. At temperatures higher than 700°C (1292°F) and in an oxidizing atmosphere, SiC starts being oxidized and is converted into a highly viscous silicate glass. This glass coating seals the pores of the refractory material very effectively. Outside of this SiC castable refractory, a layer of solidified slag is formed to protect the cooling screen surface. Over the solidified slag, molten slag flows downward by gravity, forming another flowing layer, further protecting the cooling screen from chemical and mechanical attacks from the particleladen flow, although the molten slag layer itself does also attack the solidified slag layer. Due to the high-temperature conditions of the slag, the cooling screen and the water jacket’s inner wall need to be anchored with studs. The design of the cooling screen and the membrane wall are similar. They are made of tightly spaced tubes

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Figure 1.11  Siemens cooling screen after 10 years of service (Klemmer, 2006).

connected and supported by a steel plate with a spacing approximately equal to the outer diameter of the tube. The tubes and the steel plate are welded together to form a gas-tight wall. Silicon carbide castable refractory forms a layer over the tubes and steel plate. Studs are welded on the tubes to anchor the SiC refractory. The Shell, Prenflow, and MHI gasifiers use membrane walls.

1.4.1.6  By staging feedstock and/or oxidant injections (single vs two-stage) Most of the entrained flow gasifiers inject all of the feedstock and oxidant at once. However, most notoriously, E-Gas gasifiers distribute the feedstock in two stages (Fig. 1.7). In the first stage, 75% of the feedstock (in coal-slurry form) and 100% of the oxygen are injected into the bottom cylinder by two horizontal opposing injectors to produce all of the thermal energy needed for gasification via combustion. In the second stage, the remaining 25% of the feedstock is injected by one injector without any more oxygen. The major advantage of this two-stage approach is that it confines the high-temperature combustion, with the temperature of this stage ranging from 1316°C (2400°F) to 1427°C (2600°F), to the bottom horizontal cylinder, resulting in only endothermic gasification processes (without combustion) occurring in the second stage in the vertical cylinder. This allows the gas in the second stage to remain at a relatively lowtemperature level of about 1038°C (1900°F) and, consequently, reduces the frequency of maintenance and the repair cost of the refractory bricks. The second advantage is that coal slagging only occurs in the bottom horizontal cylinder, further relieving the upper cylinder from slagging and maintenance issues. The third advantage is that using two stages allows for flexibility in adjusting the feedstock distribution between the two stages to accommodate different coal properties, especially for sub-bituminous coals, to achieve better gasification performance. However, the disadvantages are (a) the life of the refractory lining in the combustion section shortens; (b) there is relatively more unreacted char in the second stage than there is in one-stage gasifiers; and (c) the low syngas temperature in the second stage results in higher methane content (CH4), which

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induces a negative impact on carbon capture if CCS is implemented. The issue of unreacted chars can be alleviated by recycling them back to the first stage for combustion. The unreacted, dry chars are easy to burn with less oxygen, which can compensate for the higher oxygen requirement for burning slurry sub-bituminous coals. Chyou et al. (2013) employed a computational simulation of a gasification process in a cross-type entrained flow gasifier, similar to the E-Gas gasifier’s geometry. They discovered that the one horizontal injector design in the second stage could introduce a large flow circulation region in the upper vertical cylinder. This slow-moving large circulation region squeezed the through-flow to a narrow passage, which accelerated the flow and reduced the residence time of this portion of the flow, resulting in more unreacted char exiting the gasifier. Thus, they recommended adding one or two injectors to increase flow mixing, achieving more uniform flow and longer residence times in the second stage of the E-Gas type gasifier. The Mitsubishi Hitachi Power Systems (MHPS) gasifier also employs a two-stage feedstock injection scheme. Different from the E-Gas gasifier, the MHPS gasifier uses a dry-fed scheme with a membrane wall (Sakamoto, 2008). The E-Gas gasifier distributes the feedstock into two stages; whereas Tsinghua University (TU) in China has developed the TU gasifier (Fig. 1.7) which distributes oxygen in two stages. In the TU gasifier, 100% of the slurry feedstock is injected with part of the oxygen at the top of the gasifier, and the remaining oxygen is injected through opposing injectors at approximately one gasifier diameter below the main injector. The main advantage of using staged oxygen is to increase the flow mixing and distribute the high-temperature combustion process into broader regions inside the gasifier, leading to a reduction of wall temperature, reduced maintenance of the refractory wall, and increased life of the primary burner (Yue, 2013 and Zhang, 2013). In Japan, a series of dry-fed, updraft, entrained flow, two-stage gasifiers have been developed from its early laboratory demonstration scale CRIEPI gasifier, through a pilot scale EAGLE gasifier, and achieved the commercial scale MHPS gasifier. The CRIEPI gasifier was developed by the Japanese Central Research Institute of Electric Power Industry (Watanabe and Otaka, 2006) and the EAGLE gasifier was developed by the Japanese New Energy and Industrial Technology Development Organization (NEDO). The acronym, EAGLE, stands for Energy Application for Gas, Liquid, and Electricity. The CRIEPI and MHPS gasifiers are air-blown, while the EAGLE gasifier is oxygen-blown. In the CRIEPI and EAGLE gasifiers, both the oxidant and coal are fed in both stages, whereas, in the MHPS gasifier, only coal without oxidant is fed in the second stage. The EAGLE gasifier operates in oxygen-blown rich mode in the first stage at the bottom of the gasifier at around 1600°C (2912°F). In the second stage, both oxygen and coal are injected under oxygen lean (fuel rich) conditions with an operating temperature around 1150°C (2102°F) (see Fig. 1.12, Kohgami, 2007). The Industrial Technologies Research Institute (ITRI) in Taiwan has also demonstrated the advantages of employing staged oxygen gasification in a downdraft, slurry-fed, entrained flow gasifier by reducing the peak temperature in the gasifier and achieving more uniform wall temperature distribution with reduced thermal stress in the refractory brick (Wang et al., 2014). Their experiments showed that, with a split of 60% oxygen injected from the primary injector and 40% from the secondary injector (see Case 6 in Fig. 1.13), the wall temperature distribution changed from a wider

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Figure 1.12  Two-stage injection of both oxidant and feedstock in EAGLE gasifier (Kohgami, 2007).

range of 600–1500°C for single-stage oxygen injection (Case 1 in Fig. 1.13) to a narrower range of 900–1200°C for two-stage oxygen injection, with a 300°C reduction in peak temperature. Guided by the experimental results of wall temperature measurements, a computational simulation study was conducted by Wang et al. (2014) to provide more insights into the flow and temperature fields in the gasifier. The simulated gas temperature and species mole fraction distribution in the horizontal and center vertical planes in the gasifier are shown in Fig. 1.14. The gas temperature is higher at the second stage injection region than it is in the top injection region. The maximum gas temperature reaches 2600K (4220°F) at some spots at the second stage injection location and 1900K (2960°F) at the top injection stage. These temperatures are higher than in an actual gasifier because the simulation was based on the adiabatic wall boundary condition (i.e., the wall was fully insulated without heat losses). The dominant reactions near the top injection location are intense char combustion (C + ½ O2 → CO and CO + ½ O2 → CO2) and volatiles oxidation. Then, the endothermic gasification reactions occur in the region between the first and second stage injection locations, as can be seen by the increased CO and H2 concentrations and reduced temperature. The WGS also occurs, evidenced by the reduced H2O concentration. At the second stage injection location, unburned char, CO, and H2 quickly react with injected oxygen from two sides, and release the second wave of combustion/oxidation heat (C + ½ O2 → CO, CO + ½ O2 → CO2 and H2 + ½ O2 → H2O). The gasification reactions dominate the process after the oxygen is completely consumed in the region between the second stage injection location and the gasifier’s exit.

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Integrated Gasification Combined Cycle (IGCC) Technologies

Figure 1.13  The effect of two-stage oxygen injection distribution in the ITRI downdraft, two-stage, entrained flow gasifier wall temperature experimental data. The 60–40% split of oxygen between the primary and secondary injections achieved the most uniform wall temperature distribution (Wang et al., 2014).

Figure 1.14  Computationally simulated gas temperature and species mole fraction distributions for 60–40% split oxygen distribution (Long and Wang, 2014).

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Figure 1.15  Computationally simulated cross-sectional and mid-plane plots of velocity vectors and contours of temperature and mole fraction of CO2 for 60–40% oxygen distribution in the ITRI’s gasifier at the second injection location: (A) cross-sectional temperature contour and velocity vectors, (B) cross-sectional mole fraction of CO2, (C) mid-plane temperature contour and velocity vectors near second injection location, and (D) mid-plane contours of mole fraction of CO2 (Long and Wang, 2014).

Figs. 1.15 and 1.16 show the simulated cross-sectional and mid-plane plots of the velocity vectors, contours of temperature, and mole fraction of CO2 with both 60–40% oxygen distribution and 90–10% oxygen distribution, respectively. In the 60–40% case, the 40 oxygen flow injected by two opposing injectors at the second stage, impinge in the mid plane and block and squeeze the main flow laterally (Fig. 1.15A and C). The contours of temperature and mole fraction of CO2 show that the chemical reactions are affected by the local flow behavior, which is demonstrated by the high-temperature and CO2 concentration presented towards the

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Integrated Gasification Combined Cycle (IGCC) Technologies

Figure 1.16  Computationally simulated cross-sectional and mid-plane plots of velocity vectors, contours of temperature and mole fraction of CO2 at the second injection location for 90–10% oxygen distribution: (A) cross-sectional temperature contours and velocity vectors, (B) cross-sectional contours of mole fraction of CO2, (C) mid-plane temperature contours and velocity vectors, and (D) mole fraction of CO2 contours (Long and Wang, 2014).

top and bottom of the gasifier in Fig. 1.15C and D. In the 90–10% case, only 10% oxygen is injected at the second stage. The opposing injecting effect is not strong enough to block the main flow, so the combustion reaction only takes place below the second stage. This is shown by the contours of temperature and mole fraction of CO2 in Fig. 1.16.

1.4.1.7  By feedstock injection scheme (tangential vs opposing jets) For a multiple injector scheme, the injectors can be arranged to inject fuel and oxidant tangentially, such as in the first stage of the MHPS gasifier, or in opposing directions, such as in the Shell and ECUST gasifiers.

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Figure 1.17  Simulation of a slurry fed, two-stage entrained flow gasifier: (A) boundary conditions, (B) flow pathlines with the first stage injectors positioned horizontally, (C) flow pathlines with the first stage injectors tilted 30 degrees downward (Silaen and Wang, 2006).

Silaen and Wang (2006) specifically investigated the effect of multiple fuel injection schemes on the flow field and gasifier performance. They conducted a computational simulation of a two-stage, entrained flow gasifier, similar to the MHPS gasifier, but the feedstock was fed in a slurry. The boundary and feedstock feed conditions are shown in Fig. 1.17A. The simulation shows that the tangential injection scheme induces a cyclonic flow field, as evidenced from the helical pathlines shown in Fig. 1.17B. The cyclonic motion effectively increases the particle mixing and residence time, even though the average through-flow velocity remains the same as it would be without the cyclonic flow field. The velocity vectors on the horizontal plane in Fig. 1.18A for horizontal injections (Case 1) shows that the velocity in the center region of the combustor is very slow compared to the velocity in the region near the wall. This is due to the tangential injection characteristics of the low-level injectors. The plot on the vertical mid-plane shows that the vertical gas velocity in the combustor is slow in comparison to the flow velocity in the diffuser section. A strong upward flow near the wall and a slow downward flow in the center region are observed in the diffusion region (the divergent part of the gasifier). When the first stage injectors were directed downward at 30 degrees (Case 2), a core of slow-moving vertical flow, not seen in other cases, was observed in the center region of the diffuser and reductor sections. This very slow-moving vertical flow region could be explained as the result of higher stagnation pressure induced by the downward injection at the first stage. Due to this higher stagnation pressure, the downward flow seen in Case 1 (Fig. 1.18A) is impeded in Case 2 (Fig. 1.18B). In Case 3, the first stage injectors are tilted 30 degrees upward. A strong downward flow in the center region of the diffuser is induced by the entraining effect of the upward jets and results in a large annular secondary flow recirculation in the diffuser and top combustor sections. This large recirculation induces a radially outward secondary flow at the bottom region of the combustor with nearly zero vertical velocity. This radially outward flow seems to interfere with the inward fuel jets and

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Integrated Gasification Combined Cycle (IGCC) Technologies

Figure 1.18  Velocity vectors on the center vertical plane with the first stage injectors (A) positioned horizontally (B) tilted 30 degrees downward, (C) tilted 30 upward (Silaen and Wang, 2006).

subsequently reduces the strength of the cyclonic momentum, as evidenced by the weaker horizontal velocity shown in Fig. 1.18C, and the lengthened helical pathline pitches in Fig. 1.17C (as compared with Case 1 in Fig. 1.17B). Due to these differences in flow field and mixing effect, the gasifier’s performance in Cases 2 and 3 therefore suffer. Consequently, the results showed that the baseline case with all injectors positioned horizontally achieved the best gasifier performance with the highest syngas heating value. Operation with the lower (first stage) injectors angled 30 degrees downward yields the lowest gasifier performance and heating value among the three cases. Wang et  al. (2010 and 2011) conducted more studies on the effect of different designs of fuel injectors on the gasification performance. For example, they found that using multiple small injection holes in an injector could change the combustion pattern and reduce the thermal stress on the injector tip, resulting in an extended lifespan of the fuel injectors. The result of their study had been implemented and validated in the ITRI’s demonstration gasifier. They further investigated the top fuel injection designs in an entrained-flow coal gasifier and discovered that the design of two concentric injectors could provide better fuel-oxidant mixing and achieve higher heating value of syngas than the design of using four separate injections (Wang et al., 2011).

1.4.2  Summary of gasifiers and their performance Overall, the characteristics of the three major types of gasifiers can be summarized as: Fixed bed gasifiers—operate mostly at low-temperature, need less oxygen, have longer coal residence time, require large amounts of steam, and achieve low carbon conversion efficiency. Fluidized bed gasifiers—operate at moderate temperature, have relatively medium residence time, are suitable for coals with high ash fusion temperatures, and achieve medium to high carbon conversion efficiency.

●

●

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31

Entrained flow gasifiers—operate at high-temperature (>1220°C or 2230°F) with high reaction rates, produce molten slag, demand higher oxygen to coal ratio (i.e., stoichiometric ratio), require low amounts of steam, have very short residence times of a few seconds, and achieve high carbon conversion.

●

The performance of a gasifier can be evaluated by three indices: (a) Carbon conversion ratio (CCR) or carbon conversion efficiency (CCE), (b) cold gasification efficiency or hot gasification efficiency, and (c) resulting syngas heating values (HHV or LHV). Carbon conversion ratio (CCR) or carbon conversion efficiency (CCE), representing the mass percentage of total carbon in the gasifier feedstock (i.e., coal or biomass) converted to syngas, is defined as: M  (1.4) m exit ∑ imfi  c   Mi  Massof carbon in thesyngass = CCR or CCE = m in (mfc ) + recycled chars Massof carbon in the feedstockk =

m exit

(∑ y ) M ni i

c

/ M syn

m in (mfc ) + recycled chars

where, i

index, representing each species component

mfi

mass fraction of carbon in each species component

m exit syngas (product gas) mass flow rate at the exit of the gasifier, m in mass flow rate of the feedstock

mfc mass fraction of carbon in the feedstock, which can be obtained from the ultimate analysis. Mc carbon molecular weight (12) Mi molecular weight of each species Msyn syngas molecular weight ni number of carbon molecule in each mole of species, e.g., n=1 for CH4, n=2 for C2H2 yi molar or volume fraction of carbon in each species component.

The third term in Eq. 1.4 is based on mass fraction and the fourth term is based on volume fraction. For example, if the carbon mass fraction in the feedstock is 75%, the recycled char is 5%, and the product gas contains mainly CO and CO2 and some hydrocarbon products such as CH4 and C2H2, its CCR is calculated as:  12 12 12 24  (1.5) m exit mfCO2 + mfCO + mfCH4 + mfC2 H2   44 28 16 26  CCR or CCE(%) = × 100% m in (0.75 + 0.05)

m exit ( yCO2 + yCO + yCH4 + 2 yC2 H2 ) × Or =

m in (0.75 + 0.05)

12 (1.6) M syn

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Integrated Gasification Combined Cycle (IGCC) Technologies

Often the volume ratio, rather than the mass ratio, of each species is measured, so Eq. (1.6) can be more conveniently used. Although CCR provides information on how much carbon has been converted, it can’t really assess the value of the syngas. For IGCC systems, the only purpose of producing syngas is to have it burned to generate electricity. Under this stipulation, the performance of a gasifier can be evaluated by either the cold gasification efficiency (CGE) or the hot gasification efficiency (HGE) based on the syngas heating values (HHV). They are defined as:

m exit ∑ imfi (HHV)i Syngas heating value = m in HHVfeedstock Feedstock heating value   m exit  ∑ i(ni yi Mi )(HHV)i  / M syn   = m in HHVfeedstock

CGE =

(1.7)

(1.8) Syngas heating value + syngassensible heat Feedstock heating vaalue m exit  ∑ i(ni yi Mi )(HHV + h )i  / M syn m exit ∑ i[ mfi (HHV + h )i ]   = = m in HHVfeedstock m in HHVfeedstock

HGE =

where, Δh is the sensible heat, which is calculated based on the temperature difference at the gasifier’s exit and the ambient temperature. The energy content (mostly sensible heat) of the oxidizing agent (air or pure oxygen) is not included because the reference energy is based on the ambient conditions. The feedstock’s HHV is based on dry content. The CGE is defined based on the principle that the syngas needs to be cooled to near the ambient temperature in order to go through the gas cleanup process. Thus, the value of the syngas is only meaningful when it is evaluated at the (cold) ambient temperature. The CGE indicates that the gasifier’s performance is downgraded if too much heat is generated during the gasification process. In contrast, the HGE gives credit to the sensible heat. This is based on the principle that most of the sensible heat (about 80%) could be recovered one way or another, so it is arguably unfair to discount, totally, its contribution to the cycle efficiency. Even so, most literature only documents the CGE to benchmark the performance of a specific gasification process.

1.5  Syngas cooling The raw syngas exiting the gasifier is at a high-temperature, especially for entrained flow gasifiers. The raw syngas temperature can be as high as around 1480°C (2700°F). At this high-temperature, if this raw syngas could be burned directly in the GT, all of the heat generated in the gasifier that wasn’t consumed by the endothermic gasification

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33

reactions would have been saved and effectively used by the GT to produce power. In reality, unfortunately, this ideal situation doesn’t happen because the raw syngas is dirty and needs to be cleaned up. Two of the conventional cleanup processes must be performed at low temperatures, namely, mercury removal at ambient temperature around 30–38°C (86–100°F), and COS hydrolysis at 177–204°C (350–400°F). Hence, the raw syngas must be cooled all the way down to the ambient temperature through the syngas cooling system. How to preserve the high-grade thermal energy efficiently in the IGCC system during the syngas cooling process is a critical design consideration for achieving high overall IGCC plant thermal efficiency. The most practical approach is to integrate the syngas cooling system thermally with the HRSG to transfer the raw syngas’s thermal energy to producing high and/or medium pressure steam, which can be used to drive the steam turbines to produce power. Some low pressure steam can also be produced for other thermal loading in the system, such as for re-boilers in the acid gas removal process. This thermal integration partially contributes to the letter “I” in the acronym, IGCC. Four different syngas cooling schemes are typically employed in IGCC systems: radiant cooling, convective cooling, water quench, and recycled syngas quench. Depending on each individual design, different cooling schemes can be combined to achieve the optimum thermal performance, such as:

●

●

●

●

radiant-water quench-convective cooling, radiant-convective cooling, water quench-convective cooling, or recycled syngas quench-convective cooling.

A radiant syngas cooler can utilize the raw syngas thermal energy efficiently to produce high-pressure (HP) steam, followed by convective cooling. However, radiant syngas coolers are very large, expensive, and require additional maintenance. To reduce capital and maintenance costs, quench cooling has been chosen as a popular alternative cooling scheme by directly spraying water into the raw syngas stream, similar to the water curtain in ThyssenKrupp Uhde’s PDQ PRENFLO gasifier shown in Fig. 1.7. Typically, using a radiant syngas cooler can achieve about 4–5 percentage points better overall IGCC plant efficiency than direct water quench (Long and Wang, 2016). Other than reducing capital and maintenance costs, using direct water quench has another advantage—it increases the steam content in the syngas, which is needed anyway later in the WGS reactor if carbon capture is implemented. Different from the radiant and water quench cooling schemes, Shell gasifiers recycle part of the downstream syngas to the gasifier exit to mix with the fresh raw syngas in order to achieve the effect of gas quench cooling. This is followed by convective syngas cooling. A convective syngas cooler consists of all or part of the following three components: a superheater, a boiler, and an economizer. Typically, the cool water or lowgrade steam is inside a series of tubes, and hot syngas flows around the tubes. But, in the E-Gas process design, a fire-tube syngas cooler is used to cool the syngas from around 1038°C (1900°F) to 316°C (600°F). The fire-tube design contains hot syngas inside the relatively smaller volume of tubes, resulting in using much less high-grade steel material. Hence, notable cost savings have been harnessed for manufacturing fire-tube boilers in comparison with similar duty water-tube boilers.

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Integrated Gasification Combined Cycle (IGCC) Technologies

Regarding syngas cooling, some have referred to the operation of the second non-slagging stage in the E-Gas gasifier as “chemical cooling.” This phrase can be misleading when it is applied to the two-stage feedstock injection because the endothermic gasification processes in the second stage should be counted as part of the overall performance of the gasification process, which is partially assessed by the stoichiometric ratio (O2/C) of the entire gasifier. For a fixed carbon conversion ratio, the lower the stoichiometric ratio used, the better the gasification performance. The primary purpose of adding coal in the second stage is to utilize the heat produced in the first stage to produce fuel (i.e., syngas) via the endothermic processes occurring in the second stage. On the other hand, once the optimum gasification performance is achieved for a specific gasifier, adding more coal can indeed cool the syngas, but it is not a recommended practice because the cooling will not be caused by the endothermic gasification process, rather it will be caused by the sensi ble heat absorbed by a lot of coal that won’t be gasified, resulting in a lower carbon conversion ratio. Therefore, two-stage gasification should be considered for the practice of managing the syngas temperature distribution in the gasifier, as discussed in section 1.4.1.6, rather than treat it as an actual syngas cooling process, which should really start downstream of the gasifier exit, not in the second stage during the gasification process.

1.6  Gas cleanup system Depending on the properties of coals, pollutant species or their precursors form during the gasification process. The typical pollutants consist of particulates, carbonyl sulfide (COS), hydrogen sulfide (H2S), sulfur dioxide (SO2), ammonia (NH3), hydrogen cyanide (HCN), mercury (Hg), phosphorus (P), and other trace heavy metal elements, such as arsenic (As), selenium (Se), cadmium (Cd), and antimony (Sb). Most of the chlorine content of the feedstock is converted to hydrogen chloride (HCl) gas and some particulate-phase chlorides. Four of the major reactions forming pollutants are:

N 2 + 3H 2 → 2NH3 [R1.12]



CO + S → COS [R1.13]



S + H 2 → H 2 S [R1.14]



S + O2 → SO2 [R1.15]

There are many different ways to clean the syngas, but, typically, they all include the following processes: dry or wet particulate removal, wet scrubbing (for particulate, chlorine, and NH3 removal), hydrolysis (for conversion of COS to H2S, and HCN to NH3), activated carbon beds for mercury and volatile metals removal, and acid gas systems (using physical or chemical solvents) for sulfur removal. The traditional syngas cleanup steps are performed in a low-temperature environment, which requires the implementation of syngas cooling with an inevitable large loss of thermal efficiency. To overcome this problem, active research

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35

and development programs have been conducted to develop warm and hot gas cleanup systems. The most notable achievement has been the successful demonstration of warm gas cleanup technology by the Research Triangle Institute (RTI) in the IGCC facility at the Polk County Power Station in Tampa, Florida (Denton, 2015).

1.6.1  Particulates removal Particulates are formed in the forms of fly ash, bottom ash, and slag from minerals in the feedstock under combustion or gasification process. For a combustion process, the fly ash, which constitutes 85–90% of the overall ash, is a fine, light gray powder made up of glassy spheres from less than 1 micron to more than 100 microns in size (typically, 98% smaller than 75 microns and 70–80% smaller than 45 microns). The material has a bulk density of about 0.8–1 tonne per cubic meter and a maximum density of 1000–1400 kg/m3 (Sahu, 2010). Fly ash contains cenospheres–hollow spherical particles that have an especially low bulk density of 0.4–0.6 tonnes per cubic meter and constitutes up to 5% of the ash weight. Cenospheres are suitable for certain special industrial applications. Bottom ash, which constitutes about 10–15% of the overall ash, has an appearance similar to dark gray, coarse sand, and its particles are clusters of small granules, up to 10 mm in diameter (typically, 60–70% smaller than 2 mm and 10–20% smaller than 75 microns). It has a bulk density of about 1 tonne per cubic meter and a maximum density (modified) of 1200–1500 kg/m3. Most entrained flow coal gasifiers are designed to operate at temperatures above the ash fusion temperature, in which the ash melts and deposits along the wall, forming a slag layer. The properties of the ash produced by entrained flow gasifiers resemble those of the ash produced by combustion boilers, but the ash produced by moving and fluidized bed gasifiers possess some different properties. Particulate removal can be accomplished using either the dry or wet methods. The dry method includes utilizing cyclones or barrier filters, such as candle filters. The wet method mainly involves using a wet scrubber, which can also remove HCN, NH3, HCl, and H2S. The wet scrubber isn’t as thermally efficient since the syngas temperature could drop to below the water saturation temperature. The dry method of using a cyclone filter has the advantages of (a) being able to recycle unreacted carbons back to the gasifier to enhance the efficiency of the gasification process, (b) low pressure losses, (c) being usable in low pressure gasification systems, and (d) low capital and maintenance costs. However, since the cyclone utilizes centrifugal force to separate the particles, weaker centrifugal forces occur when the particulates are small and when the operation is at part-load. Thus, cyclones usually have the disadvantage of limited particulate size distribution for any specific system design. Hence, multiple cyclone stages or an additional downstream particulate removal system is needed to achieve complete particulate removal. The barrier filters don’t have the above limitations, but they induce a higher pressure loss from about 250 mbar (100 inch H2O or 3.625 psi) in new installations to about 500 mbar (200 inch H2O or 7.25 psi) in dirty units. Barrier filters allow for nearly complete particulate removal and are not affected by reduced operating loads.

36

Integrated Gasification Combined Cycle (IGCC) Technologies

The barrier filters commonly used have a shape like a long, slim candle, which is why they are often called candle filters. Candle filters are often made out of ceramics. However, due to the brittle nature of ceramics, sintered metal materials, such as iron aluminide, have also been developed. The operating temperature of candle filters ranges between 340°C and 500°C (644°F and 932°F). The upper-bound temperature (500°C) is designed to filter alkali compounds as solids because when the temperature is above 500°C, alkali compounds would go through the filter as vapor and cause corrosion by forming alkali hydro-sulfates in downstream components. The lower-bound operating temperature (340°C) is set to allow ammonium chloride (NH4Cl) to pass as a vapor instead of blinding the filter too quickly because NH4Cl can be removed by other means downstream. Candle filters are usually closely packed together in a vessel, for example, in 18 clusters with 42 ‘candle’ elements in each cluster (see Chapter 20—Puertollano IGCC Plant). Dirty raw syngas flows through the candle filters, and the particulates are deposited on the outside of the filter wall, typically with the fly ash serving as a substrate for deposits of other metallic compounds. These deposits are removed when they are blown off by nitrogen and/or other gases (such as recycled syngas) with a pulsation frequency from 4 to 40 times per hour, depending on the differential pressure across the filter. When the pressure drop across the filters reaches 1000 mbar (14.5 psi), they are due for replacement. The Tampa and Sarlux IGCC plants use wet scrubbers; the Buggenum and Puertollano IGCC plants employed ceramic candle filters; and the Wabash, Nakoso, and Kemper IGCC plants use metallic candle filters. Please see the corresponding case studies in Part V of this book.

1.6.2  Mercury removal Mercury (in the vapor phase) is traditionally removed by an activated carbon bed. In an IGCC system, a notable mercury removal process is the patented process developed by Eastman Chemical Company in Kingsport, Tennessee (Trapp, 2002), which is an activated, sulfur-impregnated, carbon bed adsorption system. The mercury removal efficiency decreases with increased temperature, so the entering syngas needs to be cooled down to about 38°C (100°F). In Eastman Chemical Company’s process, the syngas enters at an even lower temperature of 30°C (86°F). Due to this required low operating temperature, it is important to consider where to locate the mercury removal unit in the gas cleanup flow path in order to minimize the impact of reduced thermal efficiency. For example, it can be placed downstream of the COS hydrolysis unit but before the AGR unit. If it is placed upstream of the COS hydrolysis unit, more preheating is needed to raise the syngas temperature to the operating temperature (200°C/392°F) for the COS hydrolysis process.

1.6.3  Acid gas removal (AGR) sulfur removal The term Acid Gas Removal (AGR) is used to describe devices and systems that remove sulfur species from the syngas. In an IGCC system, sulfur species primarily

An overview of IGCC systems

37

appear in two forms: COS and H2S. The following section describes how to remove both COS and H2S.

1.6.3.1  COS hydrolysis Most of the sulfur in the fuel is converted to H2S under the oxygen-deficient conditions inside the gasifier: only a small amount of sulfur (3–10%) is converted to COS, which can be removed directly by an adsorption process. However, in IGCC systems, the amount of COS produced is minimal, and there is H2S that needs to be removed; thus, it is more economical and convenient to convert COS to H2S through the hydrolysis process:

COS + H 2 O ↔ H 2 S + CO2 (COSHydrolysis) [R1.16]

This reaction is slightly exothermic. Traditionally, without carbon capture, this process is placed downstream of a wet scrubber unit, which, as mentioned previously, removes particulates, HCN, NH3, HCl, and some H2S. It is typically operated at a temperature of around 200°C (392°F). Although the COS conversion efficiency is higher in lower temperature environments, if the temperature at the exit of the wet scrubber is below the saturation temperature of water, the forward reaction rate will become slow and the liquid water may poison the catalyst. Under this condition, the syngas needs to be preheated first to around 200°C in order to increase the reaction rate and vaporize all of the liquid. A company called Porocel claimed that a retention time of 50–75 seconds can achieve 99.5% conversion of the COS for the process they developed. The Porocel process uses an alumina-based catalyst (Hydrocel 640), which can promote the COS hydrolysis reaction without promoting the reaction of H2S and CO, which would form COS and H2 (Porocel, 2016 and US DOE, 2015b). The COS conversion process, although lightly exothermic, remains almost isothermal because the reaction heat is absorbed by the large amount of non-reacting components. The product gas from COS hydrolysis needs to be cooled to below 38°C (100°F) for the downstream mercury and sulfur removal processes. The heat is recovered to generate low pressure steam and/or preheat boiler feed water. For carbon capture cases, if the sour-shift option is selected, COS hydrolysis is performed together with the WGS process because the catalysts used for WGS also promote the forward COS hydrolysis reaction.

1.6.3.2  Sulfur removal Once COS is converted to H2S, many commercial processes traditionally utilized in the oil/gas, chemical, and petrochemical industries can be used to remove H2S. In IGCC systems, the process of removing H2S differs, depending on whether carbon dioxide is captured or not. The reason is that the solvents used for removing H2S can also remove CO2.

38

Integrated Gasification Combined Cycle (IGCC) Technologies

The traditional sulfur removal processes utilize a variety of commercial solvents, which can be categorized into three general types: chemical, physical, and hybrid solvents. Chemical solvents—Chemical solvents are in liquid form and are less sensitive to the acid gas partial pressure. Thus, chemical solvents are more adaptable to conditions of lower operating pressures. The advantage of chemical solvents is that they can deliver highly purified treated gases. Most of the popular chemical solvents are amine-based reagents, such as primary amine Monoethanolamine (MEA), secondary amine Diisopropanolamine (DIPA), and Methyldiethanolamine (MDEA). Chemical solvents are typically operated at temperatures that range from 27°C to 49°C (80–120°F) to absorb acid gases. The acid gases are removed by using a reversible chemical reaction operated at the boiling point of the solutions, generally from 104°C to 127°C (220–260°F) (US DOE, 2013). The solvent absorbs the acid gas in a column called an absorber. The solvent-captured gas mixture is then separated from the cleaned syngas, and the solvent is then regenerated, releasing the acid gases in a separate column called a stripper. The regenerated solvent is then sent back to the absorber, and the cycle continues. Physical solvents—The term “physical” indicates that the absorption of acid gas depends on the acid gas partial pressure, so physical solvents are preferable under high operating pressure with relatively higher H2S concentration. However, an effective absorption process using physical solvents sometimes requires temperatures lower than the ambient temperature; thus, refrigeration is required as needed, depending on the characteristics of solvents. One advantage of using physical solvents is that the solubility of H2S is an order of magnitude higher than that of CO2 at any given temperature, which allows selective absorption of H2S over CO2, which makes the process easier to control. For example, by placing the CO2 absorption column upstream while using a lean solvent (unloaded gas) and placing the H2S absorption downstream when the solvent is loaded with absorbed CO2 (see Fig. 1.20 in Section 1.7), both H2S and CO2 can be removed effectively in an integrated, compact system. The absorbed gases can be removed by reducing the pressure (or flashing) in multiple stages to regenerate the physical solvents. Common commercial physical solvents used for AGR include UOP’s Selexol (dimethyl ether of polyethylene glycol), Linde AG’s Rectisol (methanol), Shell’s Sulfolane (Tetrahydrothiophenedioide), and Lurgi’s Purisol (N-methyl pyrrolidone). (US DOE, 2013). Hybrid solvents—As the name implies, hybrid solvents consist of mixed chemical and physical solvents, so they combine the features of both types of solvents. Some examples of hybrid solvents are Lurgi’s Amisol (methanol mixed with MDEA or diethylamine), Shell’s Sulfinol (sulfolane mixed with DIPA or MEDA), and Snamprogetti Selefining (methanol mixed with toluene). Based on the above descriptions, it can be seen that conventional gas cleanup processes require syngas to be cooled to near or below the ambient temperature. This required cooling process exerts a heavy penalty on the overall plant’s thermal efficiency. Thermal integration of effective heat-exchanging processes within the IGCC system to recover this useful energy (exergy) has been a critical optimum design consideration to minimize this thermal efficiency penalty.

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39

1.6.4  Sulfur recovery The removed acid gas contains concentrated H2S. Elemental sulfur can be recovered using the traditional three-stage Claus process, which has a recovery efficiency of up to 95–98%. The Claus process converts H2S to elemental sulfur via the following reactions:

H 2 S + 3/2O2 ↔ H 2 O + SO2 [R1.17]



2H 2 S + SO2 ↔ 2H 2 O + 3S [R1.18]

The second reaction, the Claus reaction, is equilibrium limited, and sulfur conversion is sensitive to the reaction temperature. The overall reaction is:

3H 2 S + 3/2 O2 ↔ 3H 2 O + 3S [R1.19]

If more than 99.8% recovery efficiency is required in order to satisfy more stringent environmental regulations, a modified Claus process can be implemented with an addon tail gas treating unit (TGTU) process (US DOE, 2009). In order to convert H2S to S with a reaction [R1.18], a sufficient amount of SO2 needs to be present. Thus, a reaction [R1.17] becomes the necessary mechanism to provide the needed amount of SO2. Typically, one-third of the H2S is burned at temperatures ranging from 1093°C to 1427°C (2000–2600°F) in the furnace with oxygen from the air to give sufficient SO2 to react with the remaining H2S. Reaction [R1.17] is highly exothermic, and the reaction heat is recovered through a waste heat boiler to generate HP steam. The gas gives out heat to generate steam in the waste heat boiler until its temperature drops to the sulfur’s saturation temperature, then the sulfur vapor condenses to liquid phase in the condenser downstream of the boiler. The tail gas from the first condenser then goes through two to three stages of catalytic conversion with each stage consisting of a gas preheat section, a catalytic reactor, and a sulfur condenser. The liquid sulfur is collected and processed as either marketable elemental sulfur or sulfuric acid. The tail gas proceeds to the incinerator or for further processing in a TGTU.

1.6.5  Warm and hot gas cleanup To reduce the thermal efficiency penalty caused by the conventional solvent-based low-temperature gas cleanup technologies, comprehensive research efforts have been made to develop hot and warm gas cleanup (WGCU) technologies. Hot gas cleanup (HGCU) technologies, operating at temperatures of about 538°C (1000°F), are capable of removing particulates, sulfur, and alkali metals to low levels. At this temperature, the equipment and material handling can encounter serious problems, and it is also difficult to remove other contaminants, such as ammonia, chlorides, and mercury. To avoid these issues, efforts have been focused on warm gas cleanup (WGCU) technologies operating at temperatures of 420°C –476°C (700–800°F). In

40

Integrated Gasification Combined Cycle (IGCC) Technologies

WGCU, solid sorbents and membranes are used. With the support of the US DOE, RTI has successfully demonstrated the integration of WGCU in the commercially operating Tampa IGCC plant (Denton, 2015).

1.7  WGS application for pre-combustion CO2 capture In Section 1.3.5, the description of the WGS process was focused on inside the gasifier. In this section, the discussion is focused on the downstream application, particularly for shifting all CO to CO2 for easy CO2 capture. This process of capturing CO2 before the syngas is combusted in the GT is referred to as pre-combustion CCS. Although the WGS unit provides an essential process to adjust the CO/H2 ratio for applications like producing SNG or other chemicals, no WGS unit is needed for IGCC systems when carbon capture is not implemented. The WGS reaction equation is presented here again for the convenience of discussion. CO + H 2 O(g) ↔ CO2 + H 2

∆H°R = −41.0 MJ/kmol(Water-gas-shift) [R1.6]

Since the forward reaction is exothermic, at lower temperatures, the forward reaction is dominant, but the reaction rate is slower. In contrast, at higher temperatures, the reaction rate is faster, but the forward equilibrium constant is smaller. Hence, an actual WGS shifting process is typically designed with multiple stages to optimize the speed and hydrogen concentration within the syngas. In the first stage, the WGS reactor operates between 315°C (600°F, inlet) and 410°C (770°F, exit) to quickly achieve approximately 70% conversion of CO to CO2. In the second stage, the operating temperature is reduced to 230–260°C (446–500oF) to increase further the equilibrium constant and achieve approximately 90% conversion of CO. If higher purity CO2 is needed for removal, a third reactor can be added to operate at an even lower temperature of 150°C (300°F) to achieve 95% plus CO conversion. Fig. 1.19 illustrates a design of a three-stage WGS shift reactor proposed by Electric Power Research Institute (EPRI) (Kubec, 2007), which has a higher temperature in the third stage than that proposed by the author. If a hydrogen purity of more than 95% is desired, a pressure-swing adsorption (PSA) system can be used after the WGS reactors. A PSA process separates a specific gas species from a mixture of gases under pressure according to the species’ molecular characteristics and affinity for a correspondingly selected adsorbent material. A PSA process relies on the fact that, under high-pressure, gases tend to be attracted to some solid surfaces, or “adsorbed.” The higher the pressure, the more gas is adsorbed. When the pressure is reduced, the gas is released, or desorbed. Therefore, by alternating the pressure from high to low, the gas can be adsorbed from the incoming stream and desorbed into a separate, outgoing stream. Using two vessels, the process can alternate between them, making the PSA process continuous. The adsorbents can be regenerated. Usually, there are 6 or more beds operating in different states. A more detailed description of the WGS and PSA processes is described in Chapter 12.

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41

Figure 1.19  EPRI’s three-stage Water-Gas-Shift (WGS) reactors (Wang and Long, 2012b, 2014).

1.7.1  Sour-shift CCS Carbon capture, as described earlier in the general introduction in Section 1.2, depending on whether or not sulfur has been removed before or after the WGS process, two options are available: sour-shift and sweet-shift. When WGS occurs before the AGR, it is called “sour-shift,” while it is called “clean-shift” or “sweetshift” when AGR is performed afterwards (see Fig. 1.3). In sour-shift, sulfur is in the gas stream alongside the CO2; thus, AGR and carbon capture can be performed at the same time using the same equipment because AGR and carbon capture absorption techniques are very similar processes. Sometimes sour-shift can also be retrofitted onto IGCC plants that have amine-based absorption AGR systems installed. Furthermore, in the sour-shift process, the CO-shift process can also be used at the same time as COS hydrolysis with the same water supply, since the COS hydrolysis reaction (COS + H 2 O ↔ CO2 + H 2 S) and the water-shift reaction are very similar. With these afore-mentioned advantages, sour-shift CCS is cheaper than sweet-shift CCS. The downside to sour-shift CCS is the fact that it requires significant amounts of steam to perform the shift reactions, since many processes must occur all at once. Second, there is a lot of waste heat, since the water-gas shift reaction is exothermic in the direction of CO2. Thus, the gas stream gets hot, and must be cooled immediately, especially before the AGR column. Since sour-shift requires large amounts of steam,

42

Integrated Gasification Combined Cycle (IGCC) Technologies

Figure 1.20  In the sour-shift system, sulfur removal (AGR) and carbon capture can be performed together using the same equipment with, for example, the Selexol® process (Wang and Long, 2012b, 2014).

slurry-fed gasifiers or quench cooled syngas, with their higher steam/CO ratios, will be more advantageous to employ with sour-shift CCS. Recently, newly developed shift catalysts used in the Kemper IGCC plant can use steam/CO ratio feed lower than 2. It is estimated that each 0.1 reduction in the steam/CO ratio can increase the electricity production by up to 4 MWe in the Kemper plant, which produces 582MWe during the peak load. Fig. 1.20 shows the detailed layout of a sour-shift CCS system. A physical solvent is often used with the pre-combustion CCS process. In this simulated case, the Selexol process is used. The absorbers themselves operate in a cascade-like manner, with the lean solvent first absorbing CO2 in one absorber, and then sliding down to absorb H2S in a second absorber. Meanwhile, the syngas enters the H2S absorber first and flows counter to the solvent, arriving at the CO2 absorber as the second stop to undergo carbon capture. As previously mentioned in the sulfur removal section, when physical solvents are used, the solubility of H2S is an order of magnitude higher than that of CO2 at a given temperature, which allows selective absorption of H2S or CO2 in a more easily controlled manner. The carbon capture system makes use of two flash tanks because the CO2 absorber isn’t directly connected to a stripper column: it will be much easier and less expensive to use flashing to pull the captured CO2 out of solution, rather than use another heat source to raise the temperature, as there are no chemical bonds that need to be broken. The top flash tank strips about 70% of the capturable CO2 from the solvent, while

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43

Figure 1.21  In the sweet-shift system, carbon capture is placed downstream of WGS (Wang and Long, 2012b and 2014).

the lower tank handles the remaining 30% to achieve 90% overall CO2 removal. The captured CO2 is compressed to 2200 psi (150 bar) and can either be transported for some other purposes, such as advanced oil recovery, or sequestered within geological formations underground such as in salt domes, limestone deposits, or caverns. Since sour-shift occurs before AGR, sufficient steam may already exist within the syngas stream for slurry-fed or quench-cooled systems; thus, no additional steam is needed. For dry-fed or radiant-cooled systems, additional steam is needed.

1.7.2  Sweet-shift CCS The inset of Fig. 1.3 shows the sweet-shift process in the overall plant layout, while Fig. 1.21 shows the detailed layout of a sweet-shift CCS system. Additional steam is needed in the sweet-shift process because most of the steam in the syngas has already been condensed in the cooling stage before sulfur removal, and thus before the shifting occurs. To match the syngas pressure, HP steam is injected. The advantage of the sweet-shift process is that the CO2 capture plant is more simplistic than an equivalent sour-shift plant, so there is less risk of losing some CO2 during acid gas removal. But note that a complete WGS system must be installed upstream of the CCS unit. The CO2 capture plant for sweet-shift CCS has the same setup and design criteria as those of the CO2 capture section of the sour-shift plant, with two flash tanks (one stripping 70% of the capturable carbon dioxide, and the next handling the remaining 30% to achieve 90% CO2 capture), and a CO2 compressor raising the pressure to 2200 psi.

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Integrated Gasification Combined Cycle (IGCC) Technologies

1.7.3  Post-combustion CCS The CCS processes discussed together with WGS in the previous section are performed before the syngas is combusted in the GT. The merit of pre-combustion CCS technology lies in the fact that only the mass flow of the fuel (syngas) is processed in contrast to a post-combustion CCS scheme, in which a much larger mass flow of the exhaust gases, including a large amount of nitrogen and other gases, are processed. In addition, the exhaust gases are at atmospheric pressure, whereas the pressure of the syngas, before combustion, is close to the high-pressure of the gasifier. Therefore, in pre-combustion CCS, due to the much smaller volumetric flow rate and higher pressure, the sizes of the piping and components are much smaller, which results in a significant cost saving in comparison with post-combustion CCS. Considering the significant capital and operating costs for implementing post-combustion CCS, the most conceivable reason for using post-combustion CCS in an IGCC system is most likely associated with the uncertainty of CO2 emission regulations. As for now, most countries do not require CCS for coal-fired power plants. Even though some countries, such as China, Japan, and Korea, have required that all new coalfired power plants must be either highly efficient supercritical, ultra-supercritical, or IGCC plants, it is not clear when or if CCS will ever be required for many countries. Therefore, it is logical for the owners in most parts of the world to build non-CCS or CCS-compatible IGCC plants first and bet that CCS will not be required for a while in their corresponding countries. Another reason for considering post-combustion CCS for an IGCC system would lie in its more favorable reliability, availability, and attainability compared to a more complicated pre-combustion CCS IGCC plant. On the other hand, for pulverized coal power plants, post-combustion CCS is the only option, regardless of if they are subcritical, supercritical, or ultra-supercritical. The inset in Fig. 1.2 shows the post-combustion CCS system in the overall layout of the IGCC plant, while Fig. 1.22 shows the detailed post-combustion CCS system. Due to the presence of highly acid SOx and NOx in the GT exhaust, only chemical absorption is applicable in this case to separate CO2 from the rest of the GT exhaust by using an amine-based solvent, such as the primary amine Monoethanolamine (MEA), secondary amine Diisopropanolamine (DIPA), or Methyldiethanolamine (MDEA). The acid gases are captured in the absorber and then removed in a stripper by using a reversible chemical reaction, which is operated at the boiling point of the acid gas-solvent solution through a reboiler at around 138°C (280°F). The solvents are regenerated, and the cycle continues. Thanks to the nature of chemical absorption, a small portion of the chemical solvent, about 0.01% wt. is lost and must be replenished. A significant water and steam supply is needed for post-combustion CCS. If the post-combustion CCS system is added after a non-capture IGCC plant has been constructed, additional cooling water and steam generation mechanisms are required because the large amounts of water and steam that are needed can’t be tapped from the existing HRSG system or water source. Again, the captured CO2 will be compressed to 2200 psi for transportation. In total, the CCS system uses around 11% of the total plant power capacity.

An overview of IGCC systems

45

Figure 1.22  Post-combustion carbon capture system (Wang and Long, 2012a, 2014).

1.8  Combined cycle power island The combined cycle power island consists of three major components: the GT, the heat recovery steam generator (HRSG), and the steam turbine. Fig. 1.23 shows the combined power cycle after the syngas is cleaned. The cleaned syngas is sent to the GT to produce electric power. The exhaust of the GT is used to generate steam through the HRSG. The steam is used to drive a steam turbine (ST) to generate more electrical power. In the simulated case in Fig. 1.23, one GT is paired with one ST. For bigger plants, the power train is often constructed with two GTs paired with one ST. Since GTs are very special, high-tech products, they are designed and manufactured with specific operating conditions in mind. Selection of GTs must be made from the existing commercially available models with their specific ratings; whereas STs and HRSGs can be scaled up easily and manufactured to match the power required by the plant designers. Similarly, almost all of the other components in the rest of the IGCC’s plants—such as ASUs, gasifiers, gas cleanup systems, etc.—can be scaled and custom-made to match the flow rates and fuel qualities needed for the GTs. Therefore, all of the thermal-flow parameters are typically designed around the GT’s specifications. More detailed information about each component is described below.

1.8.1  Gas turbine systems The GT system, operated in the Brayton cycle, consists of a compressor, a combustor, and a turbine (or expander). An ideal Brayton cycle undergoes an isentropic

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Integrated Gasification Combined Cycle (IGCC) Technologies

Figure 1.23  Combined power cycle for sweet-shift CCS with additional steam supplied circled for CO-shift and for cooling of WGS unit circled (Wang and Long, 2014).

compression process, an isobaric combustion process, an isentropic expansion process to generate power, and an isobaric cooling process to dump waste heat to complete the cycle. In the heavy-frame industrial GT systems for electric power generation, the GT systems are commonly classified based on their power outputs and turbine inlet temperature (TIT) by using a sequence of numbers (such as 4, 5, 6, and 7) and letters in alphabetical order (such as D, E, F, G, H, and J classes), respectively. For example, for a 5F-class GT, the simple cycle power output is about 230 MW with a TIT of 1288°C (2350°F). For a 7G-class GT, the simple cycle output is about 330 MW and the TIT is about 1427°C (2600°F). For an ideal Brayton cycle, the efficiency can be increased by raising the pressure ratio (PR) and/or the turbine inlet temperature (TIT). This is what the GT industry has been doing continuously for many decades by overcoming the never-ending technical challenges and pushing the state-of-the-art GT technologies frontier forward with higher and higher values of PR and TIT. For example, from F-class industrial GTs in the 1980s to J-class in 2016, the PR value has increased from 14 to 27 and TIT has increased from 1288°C (2350°F) to 1600°C (2912°F). (Note: Aircraft engines have higher PR and TIT values than industrial GTs.) This persistent effort has resulted in consecutive record-breaking thermal efficiencies of NG-fired simple cycles of about 41% (LHV) and combined cycle thermal efficiencies reaching 62% (LHV) in 2016. In the simulated case, shown in Figs. 1.2 and 1.23, the Siemens SGT6-4000F is installed. The pressure ratio is 15.8 and the TIT is 1280°C (2335°F) for burning syngas, which is a bit lower than the rated 1288°C (2350°F) for burning NG. This is because the

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47

Figure 1.24  A cut-away view of an industrial gas turbine. Flow moves toward the right. http://upstatebusinessjournal.com/wp-content/uploads/2014/07/GE2.jpg.

products of syngas combustion contain more moisture, which can contribute to the overheating of turbine components and affect the rating criteria for the parts’ lives; thus, a reduction in syngas firing temperatures (compared to the NG firing) is required for using syngas as the fuel (Brdar and Jones, 2000). Fig. 1.24 shows a cut-away view of the flow path in an industrial GT. Air is first compressed through the compressor to a high pressure, which ranges from 15 bars to 27 bars. The compressed air is then mixed with the syngas fuel in the combustor and burned to generate high-temperature gases. The hot gases rush into the turbine, spinning the blades, and converting the hot thermal energy into rotating shaft power, which drives a generator to produce electric power. For a simple cycle, the exhaust gases at temperatures of 593–638°C (1100–1180°F) are dumped to the ambient and wasted. In a combined cycle, the remaining energy in the GT exhaust is recovered through the HRSG to generate steam for the bottom cycle. The very high temperatures in the combustor, which range from 1760°C to 1927°C (3200°F–3500°F), need to be tamed for two reasons: to (a) minimize NOx formation and (b) keep the gas temperature at the turbine inlet lower than the allowable metallurgical limit of the turbine blades. Each subject is explained in detail below.

1.8.1.1 NOx emissions NOx is a generic term representing different nitrogen oxide compounds. In GT combustion, the major components of the NOx formed consist of nitric oxide (NO) and nitrogen dioxide (NO2). NOx is typically formed via three mechanisms: (a) Fuel NOx—when nitrogen contained in the fuel combines with oxygen in the combustion air. (b) Prompt NOx—when hydrocarbons in the fuel break down and recombine with nitrogen in the combustion air, typically occurring in the fast-traveling flame fronts. (c) Thermal NOx—when the nitrogen and oxygen in the combustion air combine due to the intense heat of the combustion process. Thermal NOx typically occurs at

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temperatures between 1650°C and 1927°C (3000°F and 3500°F). NOx formation in the GT combustor is mostly thermal NOx, the mechanism for which becomes increasingly more dominant with increased temperature above 1650°C (3000°F). Thus, the most practical and effective means to minimize NOx formation is to reduce the combustion temperature. For traditional natural gas-fired turbines, premixed, dry, lowNOx burners have been used in conjunction with staged combustion to reduce NOx. In regions where water supply is not an issue, injecting steam into the combustion chamber has offered an easy, economical, and very effective scheme to significantly reduce NOx emissions. In IGCC applications, the availability of nitrogen in oxygenblown gasification systems can be used to dilute the combustion gases. For IGCC plants with CCS, CO2 is also available for use to dilute the combustion gases, but this practice will release part of the expensively captured CO2 to the atmosphere. If the NOx emissions are still not low enough to comply with current emission regulations after implementing any of the above schemes, the last resort is to install an expensive selective catalytic reduction (SCR) or selective non-catalytic reduction (SNCR) unit to remove the remaining NOx. In an SCR unit, a gaseous reductant is injected into the exhaust gas to convert NOx into N2 and water vapor with the presence of a catalyst. Anhydrous ammonia, aqueous ammonia, and urea are commonly used reductants. The catalyst components are typically made of oxides of vanadium, molybdenum, titanium, or tungsten. The catalyst is usually implemented by constructing a porous membrane that is inserted into the exhaust gas duct: this membrane typically has a limited lifespan and is easily plugged by containments. Operation and maintenance costs are high. Therefore, it is typically implemented as the last resort to remove NOx when other schemes can’t achieve the required result. The implementation scheme for SNCR is generally the same as that for SCR, except for the absence of the catalyst. As such, SNCR can be used as a cheaper, easier alternative, but it is not as effective at reducing NOx emissions as SCR.

1.8.1.2  Syngas and hydrogen-rich fuel combustion The primary fuel components in syngas are CO and H2. For partial CO2 capture systems, the H2 content increases, resulting in hydrogen-rich fuels. The energy content (heating value) of syngas is lower than that of natural gas, ranging from about 35% of the natural gas’s heating value for oxygen-blown systems to about 10% for airblown systems. The heating value can be expressed as either higher heating value (HHV) or lower heating value (LHV). The difference between these two values is the latent heat (2257 kJ/kg, 970.3 Btu/lb) of water at one standard atmosphere. HHV is based on all of the energy that can be released when the syngas is completely burned under stoichiometric conditions when all water vapor in the product gases condenses into liquid, while LHV doesn’t count the latent heat and is based on all water vapor in the final product remaining as such. The plant thermal efficiency will be higher when based on the LHV value of the fuel and lower when based on the HHV value. Table 1.3 shows examples of syngas compositions and gas heating values from selected plants and Table 1.4 lists the syngas fuel heating values compared with natural gas, CO, and H2 using various units.

Table 1.3  Snapshot

of IGCC Syngas Fuel Composition with a typical natural gas composition (Stiegel, 2009).

Syngas

PSI

Tampa EI Dorado

Pernis

ILVA

Schwarze Sarlux Pumpe

Fife

Exxon Valero Singapore Delaware

d

Natural Gas

H2 CO CH4 CO2 N2 + Ar H2O

24.8 39.5 1.5 9.3 2.3 22.7

37.2 46.6 0.1 13.3 2.5 0.3

35.4 45.0 0.0 17.1 2.1 0.4

34.4 35.1 0.3 30.0 0.2 —

8.6 26.2 8.2 14.0 42.5 —

61.9 26.2 6.9 2.8 1.8 —

22.7 30.6 0.2 5.6 1.1 39.8

34.4 55.4 5.1 1.6 3.1 —

44.5 35.4 0.5 17.9 1.4 0.1

32.0 49.5 0.1 15.8 2.2 0.4

33.4 42.2 0.1 17.8 5.7 0.1

trace — 93.9 14.5 48.2 0.9

242.0 9528.0

210.0 8274.0

183.0 317.0 7191.0 12,492.0

163.0 6403.0

319.0 241.0 12,568.0 9477.0

248.0 9768.0

230.4 9079.0

134.6 5304.0

250.0 121.0 0.79 N2/Steam

200.0 96.0 0.98 Steam

400.0 100.0 204.0 38.0 0.33 2.36 — Steam

392.0 100.0 200.0 38.0 0.74 0.62 Moisture H2O

350.0 177.0 1.26 Steam

570.0 299.0 0.65 H2O/N2

300.0 149.0 0.79 N2/H2O

— — 0.46 n/a

113° 4452.0

198.0 7801.0

— —

116.0 4660.0

150.0 5910.0

115.3 4543.0

134.6 5304.0

LHVa Btu/ft3 kJ/M3

209.0 253.0 8224.0 9962.0

GT Temperature °F 570.0 700.0 °C 330.0 371.0 H2/CO ratio 0.63 0.80 Diluent Steam N2

Equivalent LHVb Btu/ft3 kJ/M3 a

150.0 118.0 5910.0 4649.0

Pre-diluent, Post-diluent, Always cofired with 50% natural gas, d Confidential b c

200.0 7880.0

— —

c

—

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Integrated Gasification Combined Cycle (IGCC) Technologies

Table 1.4 

Comparison of typical heating values for syngases with methane, natural gas, CO, and H2 with different units Fuel type

kJ/Nm3

Cal/Nm3

Btu/scf

Methane

39,815 (HHV) 39,435 (LHV)

9510 (HHV) 9418 (LHV)

1011 (HHV) 55,530 (HHV) 910 (LHV) 55,000 (LHV)

23,811 (HHV) 21,433 (LHV)

Natural Gas (typical) O2-blown Syngas

~43,000 (HHV) ~10,200 (HHV) 1047 (HHV) 55,464 (HHV) ~33,500 (LHV) ~8000 (LHV) 912 (LHV) 49,995 (LHV)

23,845 (HHV) 21,494 (LHV)

9000 ~ 13,000 (LHV)

2150 ~ 3105 (LHV)

242 ~ 350 (LHV)

3420 (LHV)

1470 (LHV)

Air-blown Syngas

4500 ~ 5500 (LHV)

1075 ~ 1314 (LHV)

120 ~ 150 (LHV)

12,200 (LHV)

5250 (LHV)

H2

12,770 (HHV)

3050 (HHV)

325 (HHV)

141,790 (HHV) 61,084 (HHV)

10,760 (LHV)

2570 (LHV)

275 (LHV)

121,000 (LHV) 51,628 (LHV)

12,619

3014

323

10,095

a

CO

kJ/kg

Btu/lb

4368

Flame speeds: CH4=0.34 m/s (1.1 ft/s), CO= 0.55 m/s (1.8 ft/s), and H2=4.79 m/s (~15.7 ft/s) 1 Btu/ft3=37.2589 kJ/m3=8.89915 kcal/m3 1 Btu/lb= 2.326 kJ/kg N: normal condition (0°C, 1Bar or 0.98692 atm) S: Standard Condition (59°F or 60°F, 1 atm at sea level) scf: standard cubic feet a For CO, HHV and LHV are the same: no water vapor is formed during combustion.

Since the syngas heating value varies from about 35% (oxygen-blown systems) to 10% (air-blown systems) of NG’s heating value, three to eight times more fuel is needed to drive the GT in order to obtain the same rated output power as that of a natural gas power system. Consequently, the combustors—originally designed for burning natural gas—must be modified. To do so, the first step is to modify the fuel injector’s design to allow more fuel flow to pass. The second step is to make the combustor larger to accommodate the increased fuel flow rate and the dilute gas (N2 or CO2). The third step is to change the control algorithm for controlling the flow speed at the nozzle of the turbine inlet guide vane. If the nozzle cross-sectional area of the turbine inlet guide vane is too small, the flow may reach sonic speed, and the flow will be choked at off-design points. This is not an ideal condition, so the nozzle dimensions will need to be enlarged to eliminate the choking condition at part-load condition when syngas fuel is used. The fourth concern is related to the combustion stability and efficiency of the syngas fuel. Carbon monoxide and hydrogen in the syngas both have higher flame speeds than the natural gas does. Hence, consideration has to be made to avoid a potential flash-back phenomenon, which is caused by the fact that the flame propagates faster than the fuel flow speed, resulting in the flame traveling back into the fuel injectors. Flash-back is dangerous and can cause serious damage to the hardware. It has also been discovered that the conventional premixed

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51

fuel injection scheme for burning natural gas had caused detrimental resonance and humming effects when hydrogen-rich fuels were burned. Therefore, the scheme of using a diffusion flame has been adopted for burning syngas and hydrogen-rich fuels. Hydrogen fuel—When full carbon capture is implemented in an IGCC system, the syngas is converted into nearly pure hydrogen gas. Combustion of hydrogen with air produces only steam and inert N2. If oxygen only is used as the oxidant, the oxyfuel combustion product will be only the steam. In this situation, the GT becomes more like an intermediate pressure (IP) steam turbine. Hydrogen turbine systems are still in the research and development stage. No IGCC hydrogen plant has been in commercial operation yet.

1.8.1.3  Turbine blade cooling As the turbine inlet temperature continues to increase, the turbine airfoils (blades and vanes) are protected from the hot flue gas by three approaches: (a) internal cooling, (b) external film cooling, and (c) thermal barrier coating (TBC). Internal cooling is provided by passing cooling air through internal serpentine channels. Cross-stream ribs or turbulators are fabricated inside the internal channels to trigger the turbulent flow and enhance the heat transfer. External film cooling is performed by exhausting the internal cooling air to the airfoil’s outer surface through multiple rows of cooling holes drilled through the airfoils’ walls. The cooling air spreads over the airfoil’s surface, forming a layer of cooling film, which protects the airfoil’s surfaces from the hot flue gas. The cooling air for the rotors and blades is bled from the middle stage of the compressor, traveling through the shaft and the rotating disks, and finally entering the blades at the root of the blades. For vanes (or stators nozzles), the cooling air is also bled from the compressor, but is transported both through the shaft and the outer casing. Turbine airfoil cooling can consume up to 3–5% of the compressor air. A tremendous amount of research has been conducted to develop more effective and advanced turbine airfoil cooling schemes. TBCs are made of materials with very low thermal conductivity and exhibit resistance to thermal shock and thermal fatigue up to 1150°C (2102°F). A common TBC material is yttria stabilized zirconia (YSZ), which is generally deposited by plasma spraying and/or electron beam physical vapor deposition (EBPVD). TBCs can sustain very high-temperature gradients, but they are subject to spallation under extremely high thermal stresses and extended usage. To reduce the thermal stress caused by non-uniform thermal gradients in the turbine airfoils, the direction solidification method has been used to grow nickel-based superalloys into single crystal materials for turbine airfoils. With TBCs, the single crystal airfoils can sustain very high temperatures without applying external film cooling. But, they are quite expensive.

1.8.2  Air integration between GT compressor and ASU To reduce the large energy consumption of compressors used in the ASU, one approach is to take advantage of the GT compressor power by extracting part of the compressed

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Integrated Gasification Combined Cycle (IGCC) Technologies

air at the exit of the GT compressor (with pressures ranging from 14 to 23 bar or 203 to 332 psi) and sending it to the inlet of the ASU. For example, for an ASU operating at 30 bar (435 psi), an F-class GT can provide compressed air at about 15 bar, and an H-class GT can provide compressed air at about 23 bar. A boost compressor can then continuously compress the air to 30 bar. For 100% air integration, all of the air entering the ASU is provided by extracting compressed air from the GT. In addition, nitrogen from the ASU can be supplied to the GT combustor to reduce the peak combustion temperature, resulting in reduced NOx formation. This can eliminate the need to use steam injection, and hence, reduce water consumption. A fully integrated IGCC system can usually enhance the overall plant thermal efficiency by about 3 percentage points, which is significant and attractive. Li and Wang (2012) showed through a simulation of a 250 MW oxygen-blown IGCC plant with a slurry-fed gasifier operated at 31 bar (450 psi) that 50% air integration can augment the overall plant thermal efficiency about 1.67 percentage points. Both the Buggenum and Puertollano IGCC plants were fully integrated. However, a fully integrated IGCC plant is more complicated to operate and maintain, which can affect the plant availability adversely. The valuable experience obtained from the Buggenum and Puertollano plants on operating and maintaining a fully integrated IGCC plant will be given in Chapters 19 and 20, respectively.

1.8.3  Heat recovery steam generator (HRSG) HRSGs are literally heat exchangers embedded with boilers, which are particularly designed and utilized to recover waste heat. The remaining thermal energy in the GT exhaust is recovered through an HRSG to generate steam to drive steam turbines in the bottom cycle. Like a boiler, an HRSG contains evaporators with drums to boil water into saturated steam. The saturated steam is further heated in the superheater section to become superheated steam. To use the low-grade energy effectively, an HRSG also makes use of a preheating section to utilize the lowest-grade (or temperature) gas to preheat feedwater, which will undergo further heating in the economizer section. In all, an HSRG consists of four major components: the water preheater, economizers, evaporators (with drums), and superheaters. Depending on the gas flow direction, HRSGs come in two types: horizontal and vertical. In the horizontal type (see Fig. 1.25), exhaust gas flows horizontally over vertical tubes; whereas, in vertical type HRSGs, exhaust gas flows vertically over horizontal tubes. In the horizontal type, natural circulation is designed to move the water/steam in the pipes without using circulating pumps. More horizontal type of HRSGs are adopted in new plants than the vertical type. In IGCC applications, the HRSG is designed with multiple pressures, typically consisting of low pressure (LP), intermediate pressure (IP), and high-pressure (HP) sections to provide steam, which can match the pressures required for three-stage steam turbines adequately. Each pressure level will need one drum. For example, two drums are seen in Fig. 1.23 and 1.25 thus, they are two-pressure HRSGs. HRSGs can also include an extra “duct burner” section to burn more fuel to add additional energy to the steam. In a combined cycle system, this additional duct burner is a convenient and economical way to increase the total power output quickly

An overview of IGCC systems

53

Figure 1.25  A horizontal HRSG (Victory Energy, LLC, 2009).

to match peak load requirements albeit the overall thermal cycle efficiency reduces. However, in IGCC applications, a duct burner can also be used for balancing steam flow energy in case the steam integration with other parts of the plant doesn’t meet the design needs due to unexpected incidents.

1.8.4  Steam turbine systems The steam turbine system operates in a conventional Rankine cycle (as the bottom cycle), except for the fact that the boiler is replaced by the HRSG. The cycle starts with preheating the feedwater using the low-grade gas through the preheater and economizers, then the preheated water in the evaporator drums is boiled into saturated steam, and, lastly, the saturated steam is converted into superheated steam by a superheater. The conventional methods of augmenting cycle efficiency, reheating and regeneration, are usually implemented in IGCC applications. Reheating means to take partially expanded steam from the HP turbine and reheat it in the HRSG to a higher temperature similar to the steam turbine inlet temperature of the HP steam turbine. Reheating can increase the average highesttemperature in the bottom cycle and, hence, increase the overall thermal efficiency of the cycle. Reheating can also increase the quality (i.e., the water vapor content)

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Integrated Gasification Combined Cycle (IGCC) Technologies

of the exhaust steam at the exit of the LP turbine, which means that there is less liquid content in the turbine exhaust, and, hence, the blade erosion problem caused by the tiny liquid droplets in the steam can be reduced or avoided. Fig. 1.23 shows that the steam exiting the HP steam turbine is reheated to 12 bar/538°C 174.5 psi/1000°F). Regeneration means to extract part of the steam from the HP turbine and divert it to preheat the feedwater with closed and/or open feedwater heaters. In the closed feedwater heaters, the feedwater and steam do not mix: they are separated by the heat exchanger walls. In the open feedwater heaters, the steam and the feedwater are mixed. The principle of regeneration is based on the Second Law of Thermodynamics: by using low-grade steam to help preheat the feedwater instead of using high-grade energy in the boiler. In this way, the temperature difference between the heating source and the heated water is reduced, and, hence, the entropy or irreversibility generated through the heat exchanger is reduced, resulting in increased efficiency. As the steam expands in the multi-stage steam turbines, the thermal energy is converted to shaft power as the steam spins the turbine blades. The shaft power drives the generator to generate electricity. The spent steam is cooled in the condenser and condenses to liquid water, which is then preheated by low-grade steam extracted from the steam turbine before it is pumped into the HRSG to complete the cycle. The condenser is cooled by cooling water, which is the secondary coolant, and is, in turn, cooled in a separate cooling circuit by the primary coolant (typically air or water). The temperature of the primary coolant, representing the quality of the ultimate energy sink into which the waste heat of the cycle is dumped, notably affects the overall cycle output power and thermal efficiency. Common energy sinks include water sources (such as rivers, lakes, oceans, and underground water) and ambient air. Using a water resource as the energy sink is more stable than using the ambient air. However, due to the concerns of thermal pollution in natural water bodies and the conservation of groundwater, using air cooling through cooling towers, particularly the dry air cooling scheme, has shown an increasing trend.

1.8.5  Steam/thermal integration in IGCC As mentioned earlier, the existing commercialized gas cleanup processes, particularly the mercury and acid removal processes, require that the syngas be cooled to near the ambient temperature, which induces heat losses and downgrades the energy quality. These losses contribute to a notable reduction of the plant’s thermal efficiency. Therefore, steam/thermal integration in an IGCC system is an essential process to minimize these heat and heat grade losses. The major steam/thermal integrations between the power island (GT, ST, and HRSG) and the rest of the IGCC system includes the following: The syngas cooling process immediately downstream of the gasifier generates HP and IP steam for the HRSG. HP or IP steam is needed to drive the COS hydrolysis and the WGS process, and, in return, the exothermic process heat from the WGS reaction is used to heat steam.

●

●

An overview of IGCC systems

55

The cleaned syngas exiting the gas cleanup system can be preheated by the hot, raw syngas leaving the wet scrubber or COS hydrolysis reactor. This also further cools down the raw syngas, reducing the overall cooling burden. During the sulfur removal process (Fig. 1.20), steam is needed to reboil the solvent to release the acid gas. The released acid gas needs to be cooled before it is sent to the Claus plant. The CO2 compressor needs to be cooled by a large amount of water, which can be integrated with the feedwater preheating process (Fig. 1.20). The staged compressors at the exits of the ASU are often intercooled to reduce power consumption, which requires cooling water. If air integration between the GT compressor and ASU and nitrogen- dilution of combustion gas are implemented, the extracted compressed air is used to preheat the nitrogen coming out of the ASU.

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●

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●

●

1.9 Economics Economic analysis for IGCC plants is very complicated and uncertain because the technologies are not fully matured. The four commercialized IGCC plants— Tampa, Wabash, Buggenum, and Puertollano—were all started as demonstration plants. Thus, the experiences learned from these four plants are extremely valuable for advancing the technologies but might not be as helpful for the cost analysis. This is because the costs of materials and labor are constantly changing and new technologies are continuously implemented. This is compounded by current and potential future emissions policies and many other factors that can affect the international and domestic financial structures. Therefore, there have always been concerns about the uncertainty of the economic analyses associated with different kinds of IGCC plants. Furthermore, the concerns onr the extensive cost overruns of two recently constructed IGCC plants (Edwardsport and Kemper) in the United States have cast further uncertainty on whether there is sufficient information available to conduct reasonably accurate economic analyses to build IGCC plants. The cost of Duke Energy’s Edwardsport IGCC plant (CCS-ready) in Knox County, Indiana has drastically escalated from its original cost of US$1.9 billion (US$3074/kW), estimated in October 2006 to exceeding US$3.5 billion (US$5663/kW) in 2015. The cost of Mississippi Power’s Kemper County IGCC plant, which captures and sequesters 65% of its CO2 emissions, has exceeded US$6.1 billion (US$10,481/kW); compared to the original estimated cost of US$2.8 billion (US$4811/kW) at the beginning of the project in April 2010. (Note that the Kemper IGCC plant’s situation is more complicated because the cost also includes the entire onsite coal-mining business of the plant, which is not a factor for building typical power plants because coal is usually purchased under long-term contracts.) Although there could be many unforeseeable and uncontrollable factors affecting the construction costs beyond the contingencies built into the original cost analysis, relatively fundamental economic analyses isolated from those uncontrollable, complex, and often independent and unique incidents can still serve as useful and essential references. Therefore, the value of the various economic analyses presented below and in the rest of this chapter lies not in the absolute accuracy of the individual case results but in their relative values based on the fact that all cases are evaluated under

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Integrated Gasification Combined Cycle (IGCC) Technologies

the same set of technical and economic assumptions. This consistency of approach allows for meaningful comparisons among the cases evaluated.

1.9.1  Examples of economic analysis from the U.S. Department of Energy Among many economic analyses for IGCC plants, the following 2013 report from the U.S. Department of Energy provides the most comprehensive and thorough documentation in the public domain: “Cost and Performance Baseline for Fossil Energy Plants Volume 1: Bituminous Coal and Natural Gas to Electricity,” Revision 2a, Report DOE/NETL-2010/1397, September 2013. Its results are excerpted and presented below. Basically, twelve power plant configurations were analyzed, including six IGCC cases utilizing the General Electric Energy (GEE), ConocoPhillips (CoP, E-Gas, now owned by Chicago Bridge & Iron (CB&I)), and Shell Global Solutions (Shell) gasifiers, each with and without carbon dioxide (CO2) capture; four PC cases, two subcritical and two supercritical (SC), each with and without CO2 capture; and two NGCC plants with and without CO2 capture. The net power output ranges from 500 MW to 750 MW. The capacity factor (CF), defined as the fraction of the annual available day of the power plant, is 80% for the IGCC cases and 85% for the PC and NGCC cases. The CO2 capture efficiency is about 90%. The cost for coal is US$1.55/GJ ($1.64/ Million Btu) for coal (Illinois No. 6) and US$6.21/GJ ($6.55 /Million Btu) for natural gas, both on an HHV basis in 2007 USD. The 2013 analyses were further updated in 2015 (US DOE, 2015a and 2015b). The fuel prices were adjusted to US$2.78/GJ ($2.94/Million Btu) for Illinois No. 6 coal and US$5.81/GJ ($6.13/Million Btu) for natural gas, both on an HHV basis in 2011 US dollars. It is interesting to show the earlier results side by side with the updated results for comparison, as in Figures (A) and (B) below. Fig. 1.26A shows the net plant efficiency (HHV). The IGCC plant using the dry-fed Shell gasifier has the highest efficiency 42.1% (HHV) in comparison with the slurryfed GE or E-Gas gasifiers. However, the Shell case is penalized more—a reduction of about 11 percentage points of efficiency—by CO2 capture than the slurry-fed GE or E-Gas processes, which only lose 7–8 percentage points of efficiency. The penalty for employing post-combustion carbon capture is about 10–11 percentage points of efficiency reduction for sub- and super-critical PC plants and 7.4 percentage points for NGCC plants. A comparison between the 2013 and 2015 analyses (Fig. 1.26A vs. B) shows that the changes in efficiency of the identical plants are not much.

1.9.2  Capital cost and cost of electricity (COE) Economic assessments include three major results: (a) the capital cost per unit power output ($/kW), (b) the cost of electricity (COE) per unit energy output ($/kWh), and (c) the cost of capturing CO2 per unit weight of the CO2 removed ($/tonne). Since the COE is calculated over the entire useable life of the plant, it is often more commonly referred to as the levelized cost of electricity (LCOE) when COE is escalated by including an annual inflation rate over the useful lifespan of the plant.

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57

Figure 1.26  Comparison of net plant thermal efficiency (HHV) with and without CO2 capture (A) from US DOE (2013) and (B) from US DOE (2015).

Fig. 1.27A compares the capital cost ($/kW) for building different power plants with and without carbon capture in terms of total overnight cost (TOC) and total as-spent cost (TASC). The total plant cost (TPC) includes all equipment (complete with initial chemical and catalyst loadings), materials, labor (direct and indirect), engineering and construction management, and contingencies (process and project).The TOC was calculated by adding the owner’s costs to the TPC. Escalation and interest on debt

Figure 1.27  (A) Comparison of capital cost ($/kW) in terms of total as-spent cost (TASC) and total overnight cost (TOC). (Note: TOC is expressed in 2007 dollars. TASC is expressed in mixed-year 2007–2011 year dollars for coal plants and 2007–2009 mixed-year dollars for NGCC plants.) (US DOE, 2013). (B) Comparison of capital cost ($/kW) in 2011 USD (US DOE, 2015).

Figure 1.27  (Continued)

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Integrated Gasification Combined Cycle (IGCC) Technologies

incurred during the capital expenditure period were estimated and added to the TOC to calculate the TASC. (US DOE, 2013) Project contingencies were added to the engineering, procurement and construction management (EPCM) capital accounts to cover project uncertainty and the cost of any additional equipment that would result from a detailed design. The contingencies represent costs that are expected to occur. Each bare erected cost (BEC) account was evaluated against the level of estimate detail and field experience to determine the project contingency. Process contingency was added to the cost account items that were deemed to be first-of-a-kind (FOAK) or posed significant risk due to lack of operating experience (US DOE, 2013). Fig. 1.27A shows that the IGCC plant using the dry-fed Shell gasifier has the highest cost with and without CO2 capture. When carbon capture is implemented, the capital costs increase by 36–47% for IGCC plants, 75–86% for PC plants, and over 200% for NGCC plants. Without CO2 capture, IGCC plants are about 22–36% more expensive to build than PC plants; however, when CO2 capture is implemented, IGCC plants using slurry-fed gasifiers can be 5–10% cheaper to build than PC plants. NGCC plants are always the most competitive power plants with or without CO2 capture. Comparison between the 2013 and 2015 reports (Fig. 1.27A vs. B) shows that the COE increases 30–40% for all of the plants except for the NGCC plants, which have about the same COEs. Another comparison between the 2013 and 2015 reports (Fig. 1.27A vs. B) shows that the capital costs increase about 20–26% for all of the plants with or without carbon capture except for NGCC plants, where the capital costs increase by about 16%. The Shell process requires the highest capital cost to build, but it also achieves the highest plant efficiency; hence, it will be interesting to see how its COE turns out. Fig. 1.28 shows that the trend of COE is similar to that of the capital cost— in the no-capture condition, the Shell process IGCC plant still has the highest COE. PC plants and NGCC plants have similar COEs, but IGCC plants are 25–37% more expensive. When carbon capture is implemented, the COEs increase by 38–48% for IGCC plants, 80–85% for PC plants, and 46% for NGCC plants. The overall COE of the plant that uses both precombustion carbon capture and the slurry-fed GE gasifier is slightly cheaper than the PC plants, which use post-combustion carbon capture schemes. Again, NGCC plants are always the most competitive power plants when carbon capture is implemented. Yet another comparison between the 2013 and 2015 reports (Fig. 1.28A vs. B) shows that the COE increases by 30–40% for all of the plants except the NGCC plants, which have about the same COEs in both reports. Although the COEs of slurry-fed IGCC plants are comparable to those of PC plants, it is important to note that the SOx, NOx, and particulates emissions from the IGCC plants (as shown in Fig. 1.29) are significantly lower than those of the PC plants. Therefore, it will only be fair the costs of reducing PS plants’ emissions to the IGCC plants’ levels be also considered for cost comparison.

1.9.3 CO2 captured and avoided costs There are different ways to calculate the cost of removing CO2. Three common approaches are presented below.

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61

Figure 1.28  (A) Comparison of cost of electricity (COE) in mixed 2007–2011 USD (US DOE, 2013). (B) Comparison of cost of electricity (COE) in 2011 USD (US DOE, 2015).

a. The cost of CO2 captured (or removed)–This is calculated from the COE difference between analogous plants with and without CO2 capture, such as a GEE radiant cooling IGCC plant without capture versus a GEE radiant cooling plant with capture. The cost of CO2 compression is included, but the costs of CO2 transport and storage (T&S) are excluded. This is the

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Integrated Gasification Combined Cycle (IGCC) Technologies

Figure 1.29  Emission of power plants simulated by US DOE (2013).

most straightforward method of calculating the cost of capturing CO2. If the captured CO2 can be sold at a price higher than this cost, a profit can be made. This method is used for the economic analysis of cogasification of coal and biomass in the next section.



Cost of CO2 captured (or removed), (COECCS − COE NonCCS )$ / MWh (1.9) = analogous plant CO2 Captured (Tonnes / MWh )

b. The cost of CO2 avoided in reference to an analogous plant–This is the cost that will incentivize carbon capture when a carbon emission tax above this value is levied to both capture and non-capture plants using analogous technology, such as a GEE radiant cooling plant with capture versus a GEE radiant cooling plant without capture. The costs of CO2 compression, transport, and storage are included.



(COECCS withT & S − COE NonCCCS )$ / MWh Cost of CO2 avoided, = analogous plant (CO2 Emit NonCCS − CO2 EmitCCS ) (Tonnes / MWh )

(1.10)

c. The cost of CO2 avoided in reference to a specified plant (US Energy Information Administration (EIA, 2016)–This is the cost which will incentivize carbon capture when a carbon emission tax above this value is levied to both a capture and a defined non-capture reference plant, such as a GEE radiant cooling plant with capture versus a supercritical pulverized coal (SCPC) plant without capture. The costs of CO2 compression, transport, and storage are included. (COECCS withT & S − COENonCCS Supercritical PC plant )$ / MWh Cost of CO2 avoided, = with a reference plant (CO2 Emit NonCCS , SCPC − CO2 EmitCCS ) (Tonnes / MWh)

(1.11)

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63

Figure 1.30  The CO2 avoided costs in reference to corresponding analogous technology and the supercritical PC plant, respectively US DOE (2013).

The results of the US DOE (2013) analysis are shown in Fig. 1.30 using methods (b) and (c). The total average cost of CO2 avoided is US$52.9/tonne ($48/ton) for IGCC plants, US$68.3/tonne ($62/ton) for PC plants, and US$83.8/tonne ($76/ton) for NGCC plants using each corresponding analogous non-capture plant as the reference. These values have been updated in the US DOE 2015 report as US$63/tonne ($57/ton) for IGCC systems and US$71/tonne ($64.4/ton) for PC plants. Since the formulae used in the 2015 report changed somewhat from the above equations used in 2013, a direct comparison between the results of 2013 and 2015 could be confusing and misleading. Using Eq. 1.9, the CO2 capture costs were calculated in the US DOE 2015 report as US$51.1/tonne ($46.35/ton) for Shell, US$47.0/tonne ($42.6/ton) for E-Gas, and US$39.1/tonne ($35.5/ton) for both the GEE Radiant and GEE Quench processes. For NG plant, the CO2 capture cost is US$71.1/tonne ($64.6/ton). David and Herzog (2012) has also performed a CO2 capture cost analysis, which can be used for a comparison with DOE’s studies.

1.10  Cogasification of coal/biomass Biomass is a renewable energy resource. Although the combustion of biomass or biomass-derived fuels produces carbon dioxide, the net emission of carbon dioxide is zero (i.e., it is neutral CO2). Plants take in the carbon dioxide from the atmosphere

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Integrated Gasification Combined Cycle (IGCC) Technologies

via photosynthesis during their lifetime. The carbon dioxide is then returned to the atmosphere either through combustion or natural biodegrading processes. The historic splurge of cultivating energy crops to provide feedstock for producing ethanol has encountered unprecedented resistance after more careful and detailed studies showed that the outlook of cultivating energy crops to produce ethanol was not as environmentally friendly or economically attractive as previously thought. Studies have shown that cultivating energy crops to produce ethanol actually produced an unexpectedly large carbon footprint from using fossil-fuel derived fertilizers and transportation fuels for cultivating and growing the energy crops. The feedstock used also competed with food sources for both mankind and animals and had affected the existing foodsupply balance of both energy and non-energy food products in the regions where ethanol was produced. However, utilization of biomass wastes does not have the above problems, and, thus, only biomass wastes are discussed in this section. Nonetheless, there are other challenges facing the effective utilization of biomass wastes: (a) supply is limited and varies with the seasons; (b) density is low, so they are expensive for long-distance transportation; (c) they can’t be stored for extended periods of time because they rot and decay quickly (Long and Wang, 2011); (e) they contain more alkali (especially sodium and potassium) and alkaline-earth metals, which are corrosive (Bain et al., 1997); (f) they promote tar and oil formation in raw syngas; and (g) these forms of biomass could often be more expensive than coal in terms of $ per unit of energy ($/Btu or $/J). The last point is worth a special note because it has been a common misconception that biomass waste is cheaper than the biomass itself per unit weight($/tonne) or than the coal per unit energy ($/J or $/Btu). Other than the common hydrocarbons, biomass, and coal are chemically dissimilar: while coal contains more sulfur and ash, biomass typically has little ash (with few exceptions) and is almost sulfur-free. Biomass, however, contains more oxygen and nitrogen, and is more prone to producing tar. Biomass generally also has higher volatile content than coal, but coal has more fixed carbon (Loganbach et al., 2001). Due to a less complex structure, biomass is typically gasified at lower temperatures and atmospheric pressure, whereas coal requires much higher pressures and temperatures to drive out volatiles and thermally crack them to lighter products. When biomass is cogasified alongside coal, the issue of tar is no longer a concern because the high coal gasification temperature will break down tars into lighters gases.

1.10.1  Biomass pretreatment Biomass is difficult to transport through feeding equipment designed for coal because biomass cannot be easily pulverized or slurried like coal can, so continuous feeding tends to be an issue. The biggest problem in this regard is that biomass has a highly fibrous, sinewy structure, making it hard to tear up and very easy to get stuck in most machines, especially between gears and in conveyor belt drives. A few steps to alleviate this issue have been taken through several technologies: in particular, torrefaction, flash pyrolysis, and pelletization, among others (Hanssen, 2007; Bergman and Kiel, 2005; and Richards, 2008). Torrefaction is a process that converts biomass into a denser, more brittle solid form by adding heat until a temperature of 200–300°C (392–572°F) is achieved in

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65

the absence of oxygen over a period of about 5–10 minutes. During this process, the biomass releases most of its water content and loses its tough, fibrous consistency. There is an energy loss of about 10% during this process (Bergman, 2005), but the loss of mass (about 30%) is greater than the loss of energy, theoretically increasing the heating value (energy/mass) of the biomass by about 28% (Bridgeman et al., 2007). The torrefied biomass looks reddish-brown, like pellet chips, and closely resembles low-to mid-grade coals in both appearance and thermo-chemical properties (Bergman and Kiel, 2005), so is sometimes called “bio-coal.” Flash pyrolysis is performed by adding heat to about 649°C (1200°F) in less than 1 second. Flash pyrolysis converts biomass into both char and a substance called “bio-oil.” Bio-oil can be sold as a value-added product and the char can be mixed with coal to be fed into the gasifier. Pelletization is a process that compresses the woody material using a die with holes of different sizes and geometries. Under higher pressures, the lignin component in the woody material is heated and slightly plasticizes, forming a natural “glue” that holds the pellet together as it cools. Pelletization requires predrying. If the raw materials chosen have high moisture content, an energy equivalent of 18% of the resulting pellet energy content is needed to predry the materials. After predrying, the average energy required to pelletize biomass is roughly between 20 kWh/tonne and 49 kWh/tonne, which is about 5–10% of the energy content of typical wood pellets (4.7–5.2 MWh/ tonne or 7.3–8.1 Btu/lb). Pellets are commonly used for wood stoves or steam boilers to generate hot air or steam for space heating. They are also used for small capacity industrial steam applications (less than 50 MW). It is not economically attractive to feed biomass pellets alone for large utility power generation application. Most of the studies involving cofeeding biomass and coal have been conducted for combustion processes at near-atmospheric pressure in subcritical PC plants, such as those documented in papers authored by Parameswaran et al. (2009) and Jesionek et al. (2010). In IGCC plants, the biomass and coal are cofed and cogasified at highpressure instead of cocombusted in a near-atmospheric pressure environment. The Polk County IGCC plant in Tampa, Florida performed several experiments in which a wood-based eucalyptus biomass feedstock was cofed into an entrained flow GE gasifier. The results showed that the existing Coal/Petcoke-fed IGCC system was feasible for biomass, and that the emissions of NOx and SOx were reduced by about 10% (McDaniel et al., 2002). The Buggenum IGCC plant in the Netherlands also successfully cogasified biomass (50% wt.) with coal using three major biomass sources: wood, sewer sludge, and manure with about 300 tonnes per year (Kanaar, 2006). Although these experiments were technically successful, no prolonged activities of using biomass were continued due to issues associated with sustainable, reliable, and affordable supplies of biomass.

1.10.2  An example of techno-economic analysis of blending biomass in coal-fed IGCC plant Wang and Long (2012a, 2012b, and 2014) have conducted a series of technoeconomic studies to investigate the effect of cogasifying coal with 10–50 wt.% biomass on IGCC performance, CO2 emissions, and electricity cost. The baseline of

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Table 1.5 

Fuel compositions and cost of Texas lignite and sugarcane bagasse (2011 USD) Component

Texas lignite (wt.%)a

Sugarcane bagasse (wt.%)a

C H2 N2 S O2 Cl2 H2O Ash LHV (Btu/lb) Price ($/ton) $/Million Btu as received

41.3 3.053 0.623 0.7476 10.09 0 37.7 6.479 6398 19.00 1.49

43.59 5.26 0.14 0.04 38.39 0 10.39 (after pre-dry) 2.19 6714 65.00 4.84

a

GTPro internal fuel library.

the coal/biomass cogasification IGCC plants is the same plant (with 100% coal and zero biomass) introduced throughout Chapter 1 earlier. The baseline plant is fed with 4300 tons of Texas lignite in slurry form per day, and produces a net output power of 240 MWe. An oxygen-blown, entrained flow GE gasifier with quench cooling is used. CO2 is captured and compressed, but not transported. The cost analysis is based on 2011 USD. The Texas S. Hallsville lignite is cheap, costing approximately US$19/ton (EIA, 2009) or US$1.46/Million Btu (see Table 1.5). The sugarcane bagasse is used as the biomass feedstock. While sugarcane itself costs approximately US$30/ton (US Department of Agriculture, 2010), bagasse is only a portion of the sugarcane. All other by-products, like cellulose and hemicellulose, are left behind to be processed into sugar and/or ethanol. On average, about 0.2 tons of dry bagasse will be produced from one ton of sugarcane and it takes about US$10/ton to transport sugarcane bagasse. The market price of the bagasse is around US$65 per ton of wet bagasse in Louisiana (Day, 2011), which is higher than the sugarcane itself at US$30/ton. In terms of fuel energy cost, Table 1.5 shows that the cost of sugarcane bagasse is about US$4.84/Million Btu (as received), which is 3.25 times of the cost of Texas lignite. The delivered sugarcane bagasse usually contains 30–40% moisture, so the equivalent cost of dry bagasse is about US$6.92/Million Btu. The received bagasse apparently needs to be treated with predrying and torrefaction processes, which, in total, consume approximately 200 kWh of energy per ton of biomass (Long and Wang, 2014). A similar situation is likely to happen to other biomass wastes, so the above description is to emphasize that the common preconception of biomass wastes being cheaper than the biomass itself or a cheaper fuel than coal is usually not true. Plants with both post-combustion and pre-combustion carbon capture at 90% capture efficiency are studied, and the results of which are compared to those of the

An overview of IGCC systems

Table 1.6 

67

Raw syngas compositions (before gas cleaning)

Biomass/Coal ratio

0%

10%

30%

50%

CO (vol%) CO2 (vol%) CH4 (vol%) H2 (vol%) H2S (vol%) H2O (vol%) COS (vol%) N2 (vol%) LHV (Btu/lb)

14.34 9.146 0.0221 14.11 0.1575 61.43 0.0052 0.6054 1653.8

14.98 8.776 0.0274 14.76 0.1434 60.56 0.0047 0.5726 1739.8

15.47 8.726 0.0299 14.91 0.1142 60.03 0.0038 0.5374 1775.2

15.97 8.670 0.0327 15.06 0.0846 59.51 0.0029 0.5016 1811.2

baseline with no carbon capture. The plant is assumed to operate 8000 hours per year (equivalent to a capacity factor of 91.3%). The levelized cost of electricity (LCOE) is based on a plant life of 30 years with 75% of the equity being subject to straight-line depreciation. Other costs include interest, taxes, operation and maintenance (O&M) costs, and fuel costs. LCOE can also be regarded as the minimum price at which electricity must be sold in order to break-even over the lifetime of the project. The “CO2 removal cost” is calculated from the cost difference between the non-capture case and the captured cases using Eq. (1.9). When biomass is involved, all biomass feedstock is deemed completely carbonneutral. This allows for the calculation of the so-called “effective” CO2 output, which is obtained by taking the total CO2 and subtracting from it the biomass’s neutral CO2. Therefore, when 90% carbon capture is performed, the plant achieves negative CO2 emissions, which can be sold for carbon credits or used to offset the CO2 emissions of other plants from the same company. Again, the value of the economic analysis of this example lies not in the absolute accuracy of the individual case results, but in their relative values based on the fact that all cases are evaluated under the same set of technical and economic assumptions. This consistency of approach allows for meaningful comparisons among the cases evaluated. The results of this economic analysis are for commercialized Nth-of-a-kind units.

1.10.2.1  Results of coal/biomass cogasification study Blending biomass with coal affects the raw syngas composition, and, under this studied feedstock and condition, sugarcane bagasse has actually increased the syngas heating value, as shown in Table 1.6. The results for the power output and plant efficiency are shown in Table 1.7 and Fig. 1.31. Blending biomass with coal generally has a minor net positive effect on the efficiency, which reaches a maximum at 10% biomass ratio (BMR). Beyond 10% BMR, however, the efficiency begins to decrease, but it is still a bit higher than the baseline case without biomass (Table 1.7). This is due to the additional energy costs required to pretreat the biomass, as it requires much more energy (about 200 kW/ton) than raw coal (about 40 kW/ton)

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Table 1.7 

Effect of biomass/coal ratio on IGCC power output and efficiency (LHV) (Wang and Long, 2014) Biomass/Coal ratio

0%

10%

30%

50%

200,019 89,477 53,499 235,997 43.01 35.06

200,018 89,790 52,451 237,356 43.59 35.70

200,017 90,191 55,913 234,296 43.96 35.49

200,017 90,551 59,277 231,291 44.31 35.27

84,409 185,934 40.16 27.62

82,668 190,260 41.06 28.62

86,098 187,369 41.42 28.38

89,434 184,610 41.79 28.15

80,258 218,279 41.43 30.29

78,444 220,712 42.05 31.03

82,101 217,639 42.41 30.79

85,670 214,643 42.76 30.56

80,404 198,120 38.42 27.33

78,586 200,290 38.98 27.99

82,310 196,906 39.27 27.70

85,934 193,598 39.56 27.40

Baseline plant (No CCS) Gross GT Power (kW) Gross ST Power (kW) Aux. Losses (kW) Net Power (kW) Gross Eff. (LHV%) Net Eff. (LHV%) Post-combustion CCS Aux. Losses (kW) Net Power (kW) Gross Eff. (LHV%) Net Eff. (LHV%) Sour-Shift CCS Aux. Losses (kW) Net Power (kW) Gross Eff. (LHV%) Net Eff. (LHV%) Sweet-Shift CCS Aux. Losses (kW) Net Power (kW) Gross Eff. (LHV%) Net Eff. (LHV%)

Figure 1.31  Effects of biomass/coal ratio on (A) net power output and (B) thermal efficiency of IGCC plants with and without different carbon capture processes (Wang and Long, 2014).

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69

to pretreat. This added energy consumption is categorized as a part of the “auxiliary losses” in the tables, so the gross efficiency is unaffected by this, and, in fact, it continues to increase since more gross power is generated. Post-combustion CCS has a clearly negative impact on overall plant performance: 17–19 MW of steam power is lost, and the total net efficiency is reduced by nearly 7.5% in all cases, as shown in Table 1.7 and Fig. 1.31. The sweet-shift appears to be the least efficient form of CCS in the present study. While the net power is still above that of post-combustion CCS, the net efficiency unexpectedly drops to a low level, which is 0.5 percentage points lower on average than even the post-combustion cases. This is due to the fact that, unlike sour-shift, sweet-shift requires additional steam input from the steam cycle directly, resulting in an undesirable large reduction of steam turbine power output. Table 1.8 and Fig. 1.32 show the emissions data for the baseline cases. Note that the overall emissions in terms of total amount (tons/year) for each type of pollutant in Table 1.8 universally decrease with the amount of biomass added. However, on a per unit output energy basis (lbs/MWh), the CO2 emissions reach a minimum at 10% BMR. The effective CO2, on the other hand, always decreases with increasing BMR, and it becomes negative if a BMR of 10% is used. Again, the effective CO2 is determined by calculating the neutral CO2 from biomass and subtracting it from the gross CO2. In the post-combustion plants, NOx and SOx emissions have been cut drastically to very low values because the post-combustion CCS process uses chemical absorption, which allows for the direct removal of SOx and NOx. Although post-combustion CCS is expensive, it does retain one advantage over sour-shift pre-combustion CCS: more effective in removal of SOx and NOx. Table 1.9 and Fig. 1.33 show the comparison of the electricity costs ($/kWh) and the capital costs ($/kW) of the studied plants. LCOE is shown to decrease when 10% biomass is used. This is due to the increased syngas heating value. However, LCOE rises again beyond 30% due to the added extra cost of the biomass. An amount of 10% BMR appears to be the most optimal ratio, as it harnesses the lowest overall LCOE. Employing post-combustion CCS adversely adds an additional 6–7 cents/ kWh in LCOE and nearly US$500/kW in capital costs. In comparison, the CCS cost for pre-combustion CCS is about 11–12 cents/kWh. In this example, sour-shift CCS is the most economical form of CCS under the conditions set forth in this study, as it is US$500/kW cheaper in capital cost than sweet-shift and US$2500/kW cheaper than post-combustion; and about 1 cent/kWh cheaper in COE than sweet-shift and 5–6 cents/kWh cheaper than post-combustion. In other words, post-combustion CCS is the most expensive practice: about 3.7 times more costly than sour-shift CCS to remove per ton of CO2. The CO2 removal cost can be referred to as the “break-even cost” for avoiding the carbon tax. Only when the carbon tax is higher than the CO2 removal cost is implementing CCS justified. As seen in Table 1.9, a US$70/ton carbon tax would only work for those plants that include biomass in the feedstock. For pure coal, a tax of about US$72/ton is necessary to justify post-combustion CCS, whereas all of the biomass cases are below US$68/ton, with one case as low as US$63/ton. The CO2 removal cost for sweet-shift is around US$30/ton, while for sour-shift it is even lower: around

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Table 1.8 

Effect of biomass/coal ratio on IGCC emissions. All of the units are in lbs/MWh except the first top section, which is in tons/year (Wang and Long, 2014) Biomass/Coal ratio

0%

10%

30%

50%

232.5 1869.6 2,045,916

232.1 1457.7 2,042,789

231.8 1057.3 2,039,757

0.245 1.710 2155 1922

0.248 1.555 2180 1482

0.250 1.143 2205

6.58 × 10−9 0.012 262 −28

6.68 × 10−9 0.010 268 −607

6.78 × 10−9 0.007 269 −1189

0.208 2.27 2600 −8

0.211 1.79 268 −536

0.214 1.32 276

0.231 2.51 309 12

0.235 1.99 315 −579

0.239 1.47 320 −1175

Baseline plant (No CCS) (tons/year) NOx SOx Gross CO2

234.7 2157.5 2,110,246

Baseline plant (No CCS) (lbs/MWh) NOx SOx Gross CO2 Effective CO2

0.249 2.285 2235 2236

1043

Post-combustion CCS (lbs/MWh) NOx SOx Gross CO2 Effective CO2

6.73 × 10−9 0.015 277 277

Sour-Shift CCS (lbs/MWh) NOx SOx Gross CO2 Effective CO2

0.211 2.65 268 268

−1064

Sweet-Shift CCS (lbs/MWh) NOx SOx Gross CO2 Effective CO2

0.234 2.93 321 1321

US$20/ton. Note that the CO2 removal cost is based on Eq. (1.9): no CO2 transport or storage cost is included. In summary, this example of blending sugarcane bagasse with Texas lignite shows that the costs of biomass wastes ($/Btu) are often higher than those of most coals. In the current example, the most optimum condition is blending 10% (wt.) biomass, which produces the highest plant thermal efficiency, lowest LCOE, and lowest CO2 removal cost. Sour-shift appears to be superior both economically and thermally to sweet-shift CCS in the current study. (Note: this result might not be true in other

Figure 1.32  Effects of biomass/coal ratio on effective CO2 emission (with 90% CO2 capture) of IGCC plants with different carbon capture processes (Wang and Long, 2014). Table 1.9 

Effect of biomass/coal ratio on IGCC economics (2011 USD) (Wang and Long, 2014) Biomass/Coal ratio

0%

10%

30%

50%

1029.8 4363 0.1008

926.74 3904 0.0979

911.62 3891 0.1084

897.44 3880 0.1190

1490.2 8015 0.1713 0.0705 71.99

1374.8 7226 0.1631 0.0652 66.86

1359.7 7257 0.1763 0.0679 65.02

1345.4 7288 0.1895 0.0705 66.12

1164.3 5334 0.1192 0.0184 18.70

1043.1 4726 0.1146 0.0167 17.31

1027.4 4721 0.1269 0.0185 18.34

1011.5 4712 0.1392 0.0202 19.17

1181.7 5964 0.1316 0.0308 32.18

1059.9 5292 0.1264 0.0285 29.85

1044.2 5303 0.1405 0.0321 31.15

1028.2 5311 0.1547 0.0357 32.19

Baseline plant (No CCS) Capital cost ($Million) Capital cost ($/kW) LCOE ($/kWh) Post-combustion CCS Capital cost ($Million) Capital cost ($/kW) LCOE ($/kW-hr) CCS cost ($/kW-hr) CO2 removal cost ($/ton) Sour-shift CCS Capital cost ($Million) Capital cost ($/kW) LCOE ($/kW-hr) CCS cost ($/kW-hr) CO2 removal cost ($/ton) Sweet-shift CCS Capital cost ($Million) Capital cost ($/kW) LCOE ($/kW-hr) CCS cost ($/kW-hr) CO2 removal cost ($/ton)

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Integrated Gasification Combined Cycle (IGCC) Technologies

Figure 1.33  Effects of biomass/coal ratio on (A) capital cost and (B) cost of electricity for IGCC plants with and without different carbon capture processes (Wang and Long, 2014).

design arrangements). Anyone intending to add more biomass should recognize that the main advantage of using more than 10% biomass waste is not for reducing fuel cost but to take advantage of biomass’s cost-effective carbon-neutral feature because it is cheaper to use biomass waste to reduce CO2 emissions than it is to capture and sequester an equal amount of CO2.

1.11 Polygeneration As mentioned in the beginning of this chapter, the generation of electricity is only one useful function of syngas. Syngas can be used to produce liquid fuels or chemicals as value-added end products. The key concept is to adjust the CO ratio by using the WGS process. For example, with H2/CO = 2, the syngas can undergo the Fischer-Tropsch process in order to synthesize gasoline or diesel, which possess higher-energy density than syngas and are easy to store. This type of coal-to-liquid (CTL) process has been more popularly practiced in China: for example, the Yankuang Cathay polygeneration facility in Shandong province, China, has been producing 240,000 tonnes of methanol per year and an additional syngas-based net power output of 64 MWe (Zhou, 2013). The mixture of CO and H2 in syngas can also be separated into individual streams. Both CO and H2 are valuable feedstock or additives for producing chemicals. CO can also be used for the iron ore reduction process in the steel industry to provide an economical way to remove oxygen from the iron ore and leave only the iron. Table 1.10 shows the added value of various products from 1 tonne of subbituminous coal. Ammonia proves to be the most valuable end product, equivalent to US$485 for one tonne of coal. In the US, the Summit Power Group has developed the Texas Clean Energy project to coproduce power, fertilizers, and CO2 (Mattes, 2014). SCS Energy is developing the Hydrogen Energy California project, the plan for which is to coproduce power, fertilizers, and CO2 for urea production and enhanced oil recovery (EOR) (Tennant, 2015). Both of these two projects are currently facing an uncertain future. In India, Reliance

An overview of IGCC systems

Table 1.10  Added

2013)

73

values from coal-derived end products (Siemens,

Gasification

1 tonne coal (subbituminous)

1600 NM3 Syngas (CO + H2)

Power & Steam

420 NM3 SNG (@US$36/Million Btu) 2.4 MWhel power (@US$60 /MWh) 420 NM3 SNG (@ US$15/Million Btu) 0.19 tonne polypropylene (@US$1580 /tonne) 0.72 tonne methanol (@ US$450/tonne) 95 gal gasoline (@US$3.5/gal) 0.98 tonne ammonia (@US$450/tonne)

US$50

Diesel/Jet/Gas Fuels

Synthesis Gas

US$145 US$225 US$300 US$325 US$330 US$485

Iron Reduction Fuel/Town Gas

Gasification

Syngas to Liquids (FischerTropsch)

Alcohols

Revenue / tonne of coal

Carbon Source

Synthetic Natural Gas Naphtha & Waxes

Product/Volume/ Prices*

H2

Ammonia Dimethyl Ether

Methanol

Methyl Acetate

Ethylene & Propylene

Acetic Acid VAM PVA

Ketene

Acetate Esters

Diketene & Derivatives

Oxo Chemicals Acetic Anhydride

Polyolefins

Figure 1.34  Products that can be derived from gasifying coal or any carbon source (Tam, 2015).

74

Integrated Gasification Combined Cycle (IGCC) Technologies

Industries has extended a coal gasification polygeneration complex in Jamnagar that is fed with a blend of petcoke and coal to produce syngas for power generation, hydrogen, and chemical production (Maritra and Francis, 2014). More detailed discussions of polygeneration are presented in Chapter 25 by Christian Wolfersdorf.

1.12 Conclusion Irrespective of new discoveries of natural gas reserves and new techniques (such as hydraulic fracturing or fracking) being developed to increase cheaper natural gas production, coal will continue to be a major energy source to produce electricity in the world, either for economic reasons or as a strategy to safeguard national energy security and independence. The conventional way of burning coal is environmentally unfriendly; therefore, it is essential that cleaner methods of utilizing coal be developed. IGCC is one of the most promising methods to generate electricity in a more efficient, environmentally benign manner than conventional fossil-fuel combustion plants. This chapter has provided an overall introduction of all of the key processes and components in IGCC systems. Some important points are summarized below: 1. Gasification is different from combustion. The purpose of combustion is to produce heat, whereas the purpose of gasification is to produce fuels or chemicals. Therefore, during a combustion process, the stoichiometric (or a bit lower than stoichiometirc) amount of oxidant is used to completely oxidize the feedstock and obtain the maximum thermal energy output (heat); whereas, during a gasification process, as little thermal energy as possible is intended to be used (and, thus, limited oxidant is needed) to convert the feedstock to useful fuels, preserving as much of the original fuel’s chemical energy (or heating value) as desired. 2. IGCC plants operate at high pressures, so the pre-combustion gas cleanup process is much more efficient and cost-effective because only the lower volumetric flow of the fuel gas with higher partial pressures of contaminants is cleaned before combustion, as compared to post-combustion cleanup processes, which must clean all the combusted gases at much larger flow rates and at ambient pressure. 3. Employing radiant syngas cooling can increase the cycle efficiency by 4–5 percentage points over full-quench systems, but it is more expensive. 4. The existing commercially available gas cleanup processes will need to cool the syngas to near ambient temperature to remove mercury and sulfur. Due to this necessity of cooling syngas, effective thermal integration is essential to minimize the heat losses and generation of irreversibility (entropy) during the heat recovery processes. Although warm gas cleanup has been successfully developed, it has yet to reach a fully matured commercialized status. 5. Fully integrating the GT compressor with the ASU would provide a 2–3 point efficiency advantage over a plant without air integration, but it would increase operation complexityand likely degrade the plant’s reliability, maintainability, and availability. 6. IGCC provides a great platform to perform pre-combustion carbon capture. The CO2 capture cost (not the avoided cost of CO2) is about 2–3 times cheaper than it is for postcombustion carbon capture. Sour-shift and sweet-shift CCS both have their pros and cons. In the simulated study presented in this chapter, the sour-shift CCS seems to be more costeffective than sweet-shift CCS. 7. For the sour-shift CCS process, COS hydrolysis and WGS can be performed together, while AGR and CO2 capture can be performed sequentially in an integrated unit.

An overview of IGCC systems

75

8. In a simulated example, the most optimum ratio for blending biomass (e.g., sugarcane bagasse) with lignite is 10% (wt.), which produces the highest increase of plant thermal efficiency, the lowest LCOE, and the lowest CO2 removal cost. Anyone intending to add more biomass is advised to recognize that the main advantage of using more than 10% biomass waste is not for reducing fuel cost but to take advantage of biomass’ cost-effective carbon-neutral nature because it is cheaper to use biomass waste to reduce CO2 emissions than it is to capture and sequester an equal amount of CO2. 9. Continuously improved efficiency of recently developed advanced GT systems (F, G, H, J classes) will play a critical role in raising the overall cycle efficiency of IGCC. The heating value of oxygen-blown syngas is about one-third of that of NG, while that of air-blown syngas is about one-eighth of that of NG. Due to these reduced heating values, NG-fired GT fuel injectors, combustors and combustion methods need to be modified to accommodate increased syngas flow rates as well as faster CO and H2 flame speeds in order to avoid flash-back and acoustic waves caused by humming and vibration in the combustor. For example, the premixed combustion method can be replaced by diffusion flame combustion approach, and the throat area of the turbine inlet guide vanes need to be opened wider. 10. Depending on the system designs, CCS would downgrade the thermal efficiency of IGCC by 7–11 points; increase the capital costs by 36–47% for IGCC plants, 75–86% for PC plants, and over 200% for NGCC plants; increase the COE by 38–48% for IGCC plants, 80–85% for PC plants, and 46% for NGCC plants.

Nomenclatures and acronyms AGR ASU BMR CCS CRIEPI COE EPRI GHG GT HGCU HP HRSG IGCC IP ITRI LCOE LP MHPS NEDO NG O&M PR scf

Acid Gas Removal Air Separation Unit Biomass/Coal Ratio Carbon Capture and Sequestration Central Research Institute of Electric Power Industry Cost of Electricity Electric Power Research Institute Greenhouse Gases Gas Turbine Hot Gas Cleanup High-Pressure Heat Recovery Steam Generator Integrated Gasification Combined Cycle Intermediate Pressure Industrial Technologies Research Institute Levelized Cost of Electricity Low Pressure Mitsubishi Hitachi Power Systems New Energy and Industrial Technology Development Organization Natural Gas Operation and Maintenance Pressure Ratio of Gas Turbine Compressor Standard cubic feet

76

SNG ST TASC TIT TOC Tonne Ton WGCU WGS

Integrated Gasification Combined Cycle (IGCC) Technologies

Substitute Natural Gas Steam Turbine Total As-Spent Cost Turbine Inlet Temperature Total Overnight Cost Metric ton (1000 kg = 2204.6 lb) US short ton (2000 lb) Warm Gas Cleanup Water-Gas Shift

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Biography

Professor Ting Wang is currently the Director of Energy Conversion and Conservation Center and Jack & Reba Matthey Endowed Chair for Energy Research at the University of New Orleans (UNO). He is also a Professor in the Department of Mechanical Engineering. He came to UNO in 1999. Prior to UNO, he served as the director of Gas Turbine Laboratory and taught for 15 years at Clemson University. He has been involved in energy conservation and power generation in full spectrum for the past 36 years. He is an experimentalist with significant computational fluid dynamics (CFD) experience. He specializes in integrated gasification combined cycle (IGCC), gasification simulation, poly-generation, gas turbine power generation, alternative fuel generation and applications, turbomachinery, energy efficiency, and general thermal-flow engineering. He has conducted both fundamental and applied research with funding from U.S. governmental agencies and various industries. Professor Wang received a Ph.D. from the University of Minnesota at Twin Cities in 1984 and an M.S. degree from the State University of New York at Buffalo in 1981 with a major in mechanical engineering. He has published over 330 research papers and reports and was the recipient of the American Society of Mechanical Society (ASME) George Westinghouse Silver Medal for his contributions to power engineering in general. He serves on the editorial board of three international journals and is a member of the Advisory Board of International Pittsburgh Coal Conference. He was appointed by former Louisiana Governor “Mike” Foster as a member of the Louisiana Comprehensive Energy Policy Advisory Commission. He is an ASME Fellow.