Pressurized fluidized bed combustion (PFBC)

Pressurized fluidized bed combustion (PFBC)

15 Pressurized fluidized bed combustion (PFBC) T. S h i m i z u, Niigata University, Japan doi: 10.1533/9780857098801.3.669 Abstract: Pressurized flu...

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15

Pressurized fluidized bed combustion (PFBC) T. S h i m i z u, Niigata University, Japan doi: 10.1533/9780857098801.3.669 Abstract: Pressurized fluidized bed combustors (PFBCs) are fluidized bed combustors (FBCs) operated at elevated pressures. The high-temperature and high-pressure flue gas from a PFBC can drive a gas turbine to obtain mechanical energy in addition to the steam turbine, thus higher efficiency can be achieved than atmospheric FBCs which have solely steam turbines. In this chapter, basic principles of combustion and pollution control under elevated pressure, operational experiences of large-scale coal-fired PFBCs, and newly developed PFBC technology for wet sludge combustion are explained. Key words: pressurized fluidized bed, sulphur dioxide, nitric oxides, nitrous oxides, sewage sludge, combustion, power generation.

15.1

Introduction

Pressurized fluidized bed combustors are FBCs operated at elevated pressures. The high-temperature and high-pressure flue gas from the combustor can drive a gas turbine to obtain mechanical energy. Furthermore, heat can be recovered from the combustor as high-temperature and high-pressure steam to drive a steam turbine. Pressurized fluidized bed combustion technology has been developed as a high efficiency coal-fired power generation technology. Combining a gas turbine with a supercritical steam cycle, their total efficiency can be higher than that of atmospheric FBCs and pulverized coal combustion, each of which has only steam cycles. PFBC technology with a supercritical steam cycle has already been commercialized up to 360 MWe in output. A new application of PFBC is wet sewage sludge combustion. Moisture in the wet sewage sludge increases the volume of the flue gas, which acts as the working fluid of the gas turbine. The surplus energy from the gas turbine is useful to produce compressed air that is utilized in wastewater treatment processing. In this chapter, coal-fired PFBC technology for power generation is explained. In addition, recent developments related to wet sludge combustion are introduced.

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15.2

Basic principles, science and technology of pressurized fluidized bed combustion (PFBC)

Some advantages of pressurized fluidized bed combustion are the following: ∑

The hot flue gas can drive a gas turbine to produce mechanical energy in addition to thermal energy (steam). ∑ With high oxygen partial pressure, the reaction rates are expected to be increased if the reaction takes place under kinetically controlled conditions. ∑ The combustor vessel has a small plan area per unit output. Figure 15.1 presents a typical schematic diagram of a pressurized bubbling fluidized bed coal combustor for power generation, equipped with gas turbine and steam cycle. A PFBC system for power generation consists of a pressure vessel housing a combustion chamber, an air compressor, a flue gas dust removal system, a flue gas turbine, and a fuel and sorbent feed device. For power generation using coal as fuel, heat recovery from the bed is necessary to maintain the desired bed temperature. For this purpose, boiler tubes are installed in the bed. Although a bubbling fluidized bed or a circulating fluidized bed can be used as a combustion chamber of a PFBC system, bubbling fluidized bed technology has been used for large-scale PFBC plants.

Bed material transportation

Reactor

Fuel and sorbent

Bed

Fly ash Steam to turbine

Dust removal*** SCR Heat (Bag filter (deNOx) recovery or EP)

Stack

Bed material storage tank*

NH3 for SNCR (deNOx) Dust removal Pressure (Cyclones** or filters) vessel

Water Start-up burner Water Compressor

Bottom ash

Air

Generator Gas turbine

15.1 Schematic diagram of coal-fired PFBC for power generation: *, the bed material tank can be installed outside of the pressure vessel; **, cyclones can be stored in the pressure vessel; ***, dust removal is not necessary if ceramic filters are used for dust removal before the gas turbine.

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A problem of atmospheric bubbling FBC is load change, but pressurized bubbling FBC can resolve this problem. For load change in a FBC, regardless of operating pressure, both the rate of heat release by combustion (i.e., fuel feed rate and air feed rate) and the heat recovery rate must be changed simultaneously. Under atmospheric pressure condition, the increase in air feed rate increases gas velocity, but the heat transfer rate to the in-bed tubes does not increase linearly with the gas velocity, thus controlling load change in an atmospheric bubbling FBC is not an easy task. In bubbling PFBCs, however, the superficial gas velocity can be maintained at 0.8–1.0 m/s irrespective of the thermal output. Instead, total pressure and the air stoichiometric ratio are changed with the load change to meet the oxygen demand to burn fuel, as shown in Fig. 15.2. Consequently, the necessary air is supplied without changing fluidizing conditions or without changing carry-over of fines. To control the heat recovery rate from the bubbling PFBC, the bed height is controlled by withdrawing/feeding bed material from/to the bed. The heat transfer coefficient from the dense bed to immersed tubes is nearly one order of magnitude higher than that from flue gas to tubes. Therefore, the overall heat transfer rate can be controlled by changing the number of tubes immersed in the dense bed and those exposed to the gas stream in the freeboard zone. For this purpose, solid storage tanks with a solid transportation system are installed. However, the effect of pressure, which changes with load, on the heat transfer coefficient plays only a minor role. The heat transfer coefficient of 380 W m–2 K–1 at a pressure of 0.3 MPa increased slightly to 460 W m–2 K–1 at 0.9 MPa (Sakata et al., 2001). Typical changes in bed height observed in large PFBCs are shown in Fig. 15.2. The bed height, which at full-load is 3.5–4.0 m, is about 2 m at 50% load. The bed height during start-up is lower than the bottom of the tube bundles and as low as about 0.6 m. The total amount of bed material is reported as about 65 tons for 71 MWe PFBC. Therefore, large solid storage tanks are necessary for load change. The solid storage tanks can be installed either in the pressure vessel or outside. The change in bed height affects the temperature of the flue gas in the freeboard (i.e., at the combustion chamber exit and the inlet of the high-temperature dust removal system). The flue gas temperature at full load is nearly equal to the bed temperature, which is maintained at 1,110–1,130 K, but the flue gas temperature is only about 1,000 K at 75% load and 920 K at 50% load (Tsuji et al., 1999) as shown in Fig. 15.3. With the change in flue gas temperature and pressure, the gas turbine output changes with the load. Turndown from 100% to 40% is possible though gross efficiency decreases from 42.5% to 35.5% (Karita 360 MWe). The rate of load change is comparable to that of conventional pulverized coal combustors: Karita 360 MWe PFBC took only 180 min in the case of warm start-up from ignition to full load (Izaki et al., 2002). The fuel and sorbent are fed into the combustor via a wet feed system or

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Fluidized bed technologies for near-zero emission combustion 5

Bed height (m)

4 3 2

71 MWe Phase 1 71 MWe Phase 2 250 MWe A bed 250 MWe B bed

1 0 30

40

50

60 70 Load (%) (a)

80

90

100

1.2

Pressure (MPa)

1.0 0.8 0.6 0.4 71 MWe Phase 1

0.2 0.0 30

250 MWe PFBC 40

50

60 70 Load (%) (b)

80

90

100

10

O2 in flue gas (%)

672

8 6 4 71 MWe Phase 1

2

71 MWe Phase 2 0 30

40

50

60 70 Load (%) (c)

80

90

100

15.2 Change in operating parameters with load change of 250 MWe PFBC, 71 MWe PFBC without fly ash recycle to the bed (Phase 1), and 71 MWe PFBC with fly ash recycle to the bed (Phase 2). (a) Bed height*; (b) operating pressure; (c) oxygen concentration in flue gas (Shimizu et al., 2001, 2002a; Sakata et al., 2001). *, Bed height of 71 MWe PFBC is estimated from bed material mass assuming linear relationship.

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Pressurized fluidized bed combustion (PFBC) 1200

673

Bed temp. of 71 MWe PFBC = 1110–1130 K

Temperature (K)

1150 1100 1050 1000 2 MWth, Freeboard at 6 m 2 MWth, Bed

950

71 MWe, Freeboard 900 1.5

2

2.5 3 3.5 Bed height (m)

4

4.5

15.3 Change in freeboard temperature with bed height of 71 MWe PFBC (Tsuji et al., 1999) and 2 MWth PFBC (Yoshioka and Takezaki, 2000).

a dry feed system. For a wet feed system, a fuel–water mixture (paste) with water content of 20–30% was prepared and then fed by pumping. For dry feed, a rock-hopper with a pneumatic transportation system was used. The fuel is crushed to smaller than about 5 mm diameter (Tsuji et al., 1999), or about 1 mm for a dry feed system (Kaneko et al., 1999), before it is fed into the combustor. Furthermore, a sorbent (limestone or dolomite) is fed to capture SO2 in the bed. The maximum size of the sorbent, 1–5 mm, is varied to attain required desulphurization efficiency and smooth fluidization. The numbers of feed nozzles for a wet feed system were 6 for a 71 MWe unit (Goto, 1998) and 10 for a 125 MWe unit (Horiuchi and Ijiri, 2000), whereas a dry feed system used 24 for an 85 MWe system (Kaneko et al., 1999). Dust removal from flue gas is necessary to prevent gas turbine blade erosion. For dust removal at high temperatures, ceramic filters and cyclones have been used. Ceramic filters, which are made of heat-resistant porous ceramics such as Al2O3, can remove the dust nearly completely. Thereby further dust removal after the gas turbine is unnecessary. However, numerous ceramic filters are required for a large-scale PFBC. Once a few ceramic filters are broken, flue gas with a high dust load might reach the gas turbine, possibly causing unexpected shut-down of the power plant. Although the dust removal efficiency of cyclones is lower than that of ceramic filters, the two largest PFBCs, Karita 360 MWe and Osaki 250 MWe, use two-stage cyclones for dust removal to protect the gas turbine. The reason is regarded as the need to avoid unexpected shut-downs because a high standard of stability is necessary for electricity supply in advanced countries. The damaged blades can be

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replaced by scheduled (periodical) maintenance if slight erosion of turbine blades happens during the operation period. For this purpose dust-resistant gas turbines for use in high dust loads of 800 mg/Nm3 have been developed and used for Karita 360 MWe (Izaki et al., 2002). To conform with dust emission regulations when cyclones are used, additional dust removal using a bag filter or electric precipitator (EP) is necessary before the flue gas can be released to the atmosphere.

15.2.1 Combustion under elevated pressure The sequence of events of fuel particles after introduction into pressurized FBCs is fundamentally identical to that in atmospheric FBCs: evaporation of moisture at first with subsequent volatile matter evolution, ignition, combustion of char, and finally entrainment of fine char/ash particles by a gas stream or withdrawal of coarse ash particles from the bottom. Details of the combustion reactions are discussed in Chapter 7. Only the effects of pressure on fuel combustion are described briefly here. With increasing pressure, the boiling point of water becomes higher, thereby delaying moisture evaporation. The boiling point under PFBC conditions is 150–180°C. Therefore, the time necessary for heating up a fuel particle to the boiling point is regarded as about twice as long as that in atmospheric FBCs if the heat transfer is the same. Additional time is necessary for heating when fuel is fed as a fuel–sorbent–water paste and large clumps of the paste are formed in the bed. Volatile matter evolution is also affected by the pressure. The yield of volatile matter is known to decrease with increasing pressure (Gadiou et al., 2002; Niksa et al., 2003). The yield of volatile matter at 1 MPa is about 10–20% less than that at atmospheric pressure. In particular, the yield of tar is reduced by elevating pressure because of inhibited evaporation of liquid at elevated pressure and simultaneous solid carbon formation within the fuel particles. Although the volatile matter yield is reduced slightly, the effect of pressure on the homogeneous reaction (oxidation) rate of volatile matter is regarded as more pronounced. Roughly speaking, collision frequency among gas molecules is proportional to the square of pressure. Consequently, the overall combustion rate, i.e., heat release rate, per unit volume of the bed is regarded as increased with elevating pressure. The number of fuel feeding points must be designed to avoid hot-spot formation near the fuel feeding point. The formation of hot-spots is determined by the balance between the rate of volatile matter evolution followed by ignition and the rate of fuel dispersion in the bed. For this purpose, wet fuel feed is regarded as advantageous because it delays the volatile matter evolution rate. Indeed, the number of fuel feeding points per unit of power output for wet fuel feed is less than that for dry fuel feed.

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The effect of pressure on the char combustion rate differs depending on the rate controlling step. In general, the gas–solid reaction rate is governed by both the chemical reaction kinetics and the diffusion of gaseous reactant(s) from the bulk gas to the solid reactant. The intrinsic reaction rate of carbon combustion (rate of carbon consumption per unit internal surface area) increases concomitantly with increasing oxygen partial pressure. To the extent that the combustion rate is governed by the chemical reaction, the combustion rate increases with increasing total pressure. However, with increasing intrinsic reaction rate, i.e., with decreasing chemical kinetic resistance, the contribution of diffusion resistance to overall resistance becomes more pronounced. Although the oxygen concentration (mole oxygen in unit volume of gas) increases with elevating total pressure, the diffusivity of oxygen in the gas is inversely proportional to the total pressure. Consequently, no influence of total pressure on the char combustion rate will be observed when the char combustion rate is governed by the boundary layer diffusion or diffusion through the ash layer formed at the external surface. Sufficiently small char particles are regarded as burnt under chemical kinetic controlling conditions with sufficient penetration of O2 into the particle’s center, i.e., with an effectiveness factor of unity. With increasing particle size, oxygen consumption near the external surface is more pronounced before oxygen diffuses towards the particle’s center, i.e., the effectiveness factor decreases. For coarse particles with a thick boundary layer, the combustion rate is governed mainly by the gas diffusion rate. Additional discussion of coal combustion under elevated pressure, including advantages and disadvantages, is presented in Section 15.4.2.

15.2.2 Basic principle of emissions control Sulphur oxides One advantage of PFBCs is in-situ SO2 capture by sorbent without external wet scrubbing system. This advantage is shared with atmospheric FBCs. Crushed natural limestone (CaCO3) or dolomite is usually used for this purpose. Furthermore, sorbent particles play a role of bed material which transfers heat from burning fuel to the immersed boiler tubes. Consequently, during combustion of low sulphur content fuel, the sorbent feed requirement is determined not for sulphur removal but to maintain sufficient bed material to ensure heat transfer. In PFBCs, the chemical form of calcium changes with operating conditions. When the partial pressure of CO2 is lower than equilibrium partial pressure for the calcination reaction, CaCO3 is decomposed to CaO (Eq. [15.1]), whereas CaCO3 is stable under higher CO2 partial pressure conditions.

CaCO3 Æ CaO + CO2

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[15.1]

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The equilibrium partial pressure of CO2 at a typical FBC temperature (1,123 K) is about 0.05 MPa. Figure 15.4 shows the CaO/(CaCO3+CaO) molar ratio of bed material drained from Osaki 250 MWe PFBC. With decreasing load, total pressure decreases. Consequently, the content of CaO increased because of the reduced partial pressure of CO2. The reaction of CaO with SO2 is fundamentally identical to that occurring in atmospheric FBCs as CaO + SO2 +

1 2

O2 Æ CaSO4.

[15.2]

At higher loads, CaCO3 is the major component of the sorbent. Direct sulphation occurs as CaCO3 + SO2 +

1 2

O2 Æ CaSO4 +CO2.

[15.3]

Because uncalcined limestone (CaCO3) is non-porous, the reaction proceeds from the external surface toward the center remaining product (CaSO 4) layer at the surface. The reaction rate of direct sulphation is therefore governed both by chemical kinetics and by diffusion of SO2 through the product layer formed at the external surface of particles. The rate of this reaction has been investigated mainly using thermogravimetric analysis (Hajaligol et al., 1988; Snow et al., 1988; Krishnan and Sotirchos, 1993a, 1993b; Zevenhoven et al., 1998; Qui and Lindqvist, 2000). The reaction rate constant, order of reaction, and effective diffusivity through the product layer have been evaluated for limestone of different types. One example of the reaction rate expression of specific reaction rate rS (kmol m–2 s–1) is r S = kSC n,

[15.4]

where C, kS and n respectively denote the concentration of SO2 (kmol m–3), sulphation rate constant, and reaction order. Qui and Lindqvist (2000) reported k = 2.08 ¥ 10–5 kmol0.42m–0.26 s–1 and n = 0.58. The effective diffusivity

CaO/(CaCO3+CaO) molar ratio (–)

1 0.8

50% load 1138 K

A boiler

100% load

0.6 0.4

75% load

1131 K

0.2 0 0.3

B boiler

0.4

0.5

1130 K

1151 K

1136 K

1132 K

0.6 0.7 Pressure (MPa)

0.8

09

1

15.4 CaO/(CaCO3+CaO) molar ratio of bed material drained from the bottom of Osaki PFBC (calculated from data reported by Hokari et al. (2001)).

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through the product layer is about 1.5 ¥ 10−5 m2 s–1 and is dependent on solid conversion. It should be emphasized that this reaction kinetics is dependent on the sorbent type and that it should be evaluated for each sorbent. Figure 15.5 shows that the reaction rate is estimated as controlled mainly by diffusion resistance when the product layer thickness is greater than 1–2 mm. The thickness of the sulphur-rich layer formed at the surface of limestone from 71 MWe PFBC was reported as 30–40 mm (Abe et al., 2000a). For that reason, it is regarded as necessary to consider diffusion resistance through the product layer. One difference between the actual PFBC system and experimental apparatus such as thermogravimetric analyzers or fixed bed reactors to measure the sulphation rate is the attrition of particles. In PFBCs, particles move in the bed and mutually collide. Therefore, the removal of the surface occurs by attrition. Results from the Wakamatsu 71 MWe PFBC demonstration project revealed that about 70% of the fed limestone was removed from the reactor as fly ash formed by attrition (Fig. 15.6). The average rate of attrition (rate of decrease in radius) was reported as about 1–2 mm/h. The attrition of the particle surface is regarded as playing an important role in determining the SO2 capture rate. One possibility is that removal of product layer at the surface by attrition reduces the diffusion resistance, thereby increasing the reaction rate. Another possibility is that the loss of limestone increases when unreacted limestone fine particles are formed and carried over by the gas stream. A reaction model with solid attrition is therefore necessary to predict desulphurization behavior or to analyze the operation results. A reaction Temp. = 1123K, P TOTAL = 1.0 MPa, De = 1.5 ¥ 10–9m2/s, r = kCsn, n = 0.58 k = 2.08 ¥ 10–5 (kmol0.42/m0.26s) 1 0.9

200 ppm 100 ppm 50 ppm

0.8 Cs /COUT (–)

0.7 0.6 0.5 0.4 0.3 0.2 0.1 0

0

1

2 3 4 5 Thickness of CaSO4 layer (µm)

6

15.5 Calculated concentration of SO2 at unreacted core surface (Cs) based on kinetic data reported by Qui and Lindqvist (2000), assuming constant diffusivity (De): COUT, concentration of SO2 at the particle surface.

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Ca attrition/Ca feed (–)

1 0.8 0.6 0.4 Without cyclone ash recycle

0.2

With cyclone ash recycle 0 20

30

40 50 60 Plant output (MWe)

70

80

15.6 Loss of limestone by attrition (calculated from results reported for the Wakamatsu 71 MWe PFBC demonstration project by Shimizu et al., 2001). SO2 conc.

Product layer thickness, d

CaSO4

COUT

CaCO3

Removed by attrition Cc Re(0) Initial surface

Re Particle surface

Rc Unreacted core surface

To particle’s center

15.7 Shrinking unreacted core model with ash layer diffusion resistance and surface attrition.

system with attrition is shown schematically in Fig. 15.7. The particle surface moves toward the center with removal of the particle surface by attrition. A product layer is formed at the surface if the reaction rate (formation rate of the product layer) is higher than the attrition rate. To establish a sulphation–attrition model of the sorbent, the mode of attrition must also be examined: either continuous attrition or intermittent attrition (Fig. 15.8). Even when the partial pressure of CO2 at the top of the bed becomes greater than the equilibrium partial pressure of calcination, the partial pressure at the bottom is regarded as lower because air without CO2 is fed from the bottom. Consequently, calcination is regarded as occurring in the CO2-lean zone at the bottom, as suggested by Abe et al. (2000a) and Hokari et al. (2001). Because limestone becomes porous with calcination, the solid strength is regarded as reduced with calcination. Aside from the continuous mechanical attrition, it is possible that chemically assisted attrition occurs with the intermittent movement of solids between the CO2-rich zone and © Woodhead Publishing Limited, 2013

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Fragments with finite thickness

CaSO4

Re

at

Rc CaCO3 Continuous decrease in size with time dRe/dt = – a Continuous attrition

Decrease in size with an interval of t Intermittent (periodical) attrition

CaO/(CaCO3 + CaO) molar ratio (–)

15.8 Continuous attrition and intermittent attrition with an interval of t for the same average attrition rate a (rate of decrease in radius). 1

Bottom (2 cyc) Prim. cyclone (2 cyc)

0.8

Bag (2 cyc) 0.6

Bottom (CTF) Cyclone (CTF)

0.4

Ceramic filter (CTF) 0.2 0 40

60

Load (%)

80

100

15.9 CaO/(CaCO3+CaO) molar ratio of ash drained from different parts of 71 MWe PFBC: 2 cyc – two-stage cyclone with bag filter at the end; CTF – ceramic tube filter after cyclone (calculated from data by Harada, 1998).

CO2-lean zone, as suggested by Abe et al. (2000a). Indeed, the composition of calcium compounds in the fly ash differs from that in the bed, as shown in Fig. 15.9. Fly ash, which contains fine calcium particles formed by attrition, was found to be rich in CaO, although the content of CaO in the bottom (bed ash) was low. This difference suggests that the fragile CaO part is removed by attrition, although hard CaCO3 remains in the bed. Model calculations incorporating intermittent nature of attrition showed that the reaction rate as well as solid utilization efficiency is affected strongly by the mode of attrition (Fig. 15.10). Even when the average attrition rate is the same, the mode of attrition is inferred to affect both the average sulphation rate and the solid utilization efficiency. For continuous attrition or intermittent attrition with a higher sulphation rate (higher SO2 concentration), the surface

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Fluidized bed technologies for near-zero emission combustion Continuous, SO2 100 ppm

250

Re, Rc (µm)

240 CaSO4

230

CaCO3

220 210

Re Rc 0

1000 Time (min) (a)

2000

10 hr period, SO2 1000 ppm

250 240 Re, Rc (µm)

230 220

CaSO4

210

CaCO3

200 Re

190 180

Rc 0

2000 Time (min) (b)

4000

10 hr period, SO2 20 ppm

250

CaSO4 240 Re, Rc (µm)

680

CaCO3 230

Removed by attrition

220 210

Re Rc 0

1000 Time (min) (c)

2000

15.10 Calculated formation of CaSO4 layer at sorbent surface under the same attrition rate condition as that in the case of (a) continuous attrition, (b) intermittent attrition under high product layer formation rate condition, and (c) intermittent attrition under slow product layer formation rate condition. Re – radius of particle; Rc – radius of unreacted core; average attrition rate a = 1 mm/h (reproduced from Shimizu et al., 2002b, with permission by Japan Institute of Energy).

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of the sorbent is invariably covered with the product layer, whereas unreacted surface appears (i.e., unreacted core is also removed by attrition) when intermittent attrition occurs at lower sulphation rate conditions (Shimizu et al., 2002a, 2002b). The criterion for appearance of an unreacted surface with intermittent attrition if the sulphation process is controlled solely by diffusion of SO2 through the product layer is as shown below:

t > 2DeCM/ra2

[15.5]

where De, M, r, t, and a respectively stand for the effective diffusivity of SO2 through the product layer (m2 s–1), molecular weight of CaCO3 (kg kmol–1), density of limestone (kg-CaCO3 m–3), interval of attrition (s), and the average attrition rate (m s–1). The overall reaction rate, as well as the reaction order with respect to the SO2 concentration, is affected by the formation of the sulphated layer at the surface. When the CaSO4 layer always exists at the surface (Figs 15.10(a) and (b)), the overall reaction rate is controlled by the attrition rate, i.e., the rate is independent of SO2 concentration. The specific SO2 capture rate per unit surface area rsteady (kmol m–2 s–1) under this condition is given as

rsteady = ar/M.

[15.6]

For this case, the product layer thickness is determined so that the moving rate of the unreacted core surface is the same as the moving rate of the particle surface. When a fresh CaCO3 surface appears at the instant of intermittent attrition (Fig. 15.10(c)), the growth rate of product layer thickness, d(m), with time is given as

dd/dt = DeCM/rd,

[15.7]

assuming that the reaction rate is controlled by diffusion through the product layer. By integrating Eq. [15.7] from t = 0 to t, the average SO2 capture rate per unit surface area under a steady condition is obtained as

rsteady = (2DerC/Mt)1/2.

[15.8]

Consequently, the apparent reaction order is 0.5 with respect to SO2 concentration. When the attrition rate is so high that unreacted CaCO3 is always exposed to the gas, then the reaction rate is governed by the chemical kinetics given in Eq. [15.4]. As described above, the dependence of the SO2 capture rate on the SO2 concentration under solid attrition conditions is weak, which suggests that the SO2 capture rate might be insufficient when burning high sulphur content fuels. Thus the SO2 concentration in the gas is anticipated to be high. This subject is discussed in Section 15.4.4. Although the mode of attrition (or interval of intermittent attrition) is an important factor, no proper experimental approach is currently available to ascertain the mode. One approach is to measure the size distribution of Ca-

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rich fly ash particles. That size is expected to be of the order of magnitude of at (attrition rate multiplied by interval of intermittent attrition). However, that investigation remains to be examined in future work. Nitrogen oxides (NOx) and nitrous oxide (N2O) The reaction mechanisms (chemical reactions) of NOx and N2O in PFBCs are fundamentally the same as those in atmospheric FBCs. As a matter of principle, the effects of elevated pressure on chemical reactions can be summarized as ∑

increased gas phase (homogeneous) reaction rates with increased partial pressure, ∑ increased gas–solid reaction rate when the reaction is controlled by reaction kinetics, and ∑ little influence on gas–solid reaction when the reaction is controlled by gas diffusion through the gas film around the particle because diffusivity of gas is inversely proportional to total pressure.

Furthermore, the change in the chemical form of sorbent (CaO or CaCO3) is regarded as affecting the emissions of NOx and N2O because of the change in chemical reactivity and porosity of the solids. Details of emissions of NOx and N2O from a large demonstration plant (Wakamatsu 71 MWe) have been reported. Tsuji et al. (1999) and Abe et al. (2001) proposed empirical equations (ASH(TR) equations) of conversions of fuel-N to NOx and N2O as shown below:

Conversion to NOx = F1(TC, PO2)/(1 + F1(TC, PO2))



– F2(TC, PO2)/(1 + F2(TC, PO2))

[15.9]



Conversion to N2O = F2(TC, PO2)/(1 + F2(TC, PO2))

[15.10]



F1(TC, PO2) = 5.00 ¥ 10–4PO21/2 exp(6.21 ¥ 103/TC)

[15.11]



F2(TC, PO2) = 4.57 ¥ 10–7PO21/2 exp (12.1 ¥ 103/TC)

[15.12]

In those equations, TC (K) and PO2 (atm) respectively denote the flue gas temperature at the cyclone (i.e., at the exit of freeboard) and partial pressure of oxygen in the flue gas. Sakuno et al. (2002) proposed a different type of empirical equation to predict NOx emissions (ppm) (corrected to dry flue gas containing 6% O2) from 71 MWe PFBC taking account of sulphur retention as:

[NOx]cal = a0 + a1([O2]–[O2]typ) + a2(hS–hS,typ) + a3(Tbed–Tbed, typ),



[15.13]

where [O2], hS, and Tbed respectively represent the oxygen concentration © Woodhead Publishing Limited, 2013

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(vol.%) in the flue gas at SCR inlet, sulphur retention efficiency (–), and bed temperature (K). The subscript typ signifies typical operation conditions. The values of [O2]typ, hS, typ, and Tbed, typ respectively represent 6.25%, 0.928 (–), and 1122 K. The coefficients ai are given by nitrogen content, N [daf%], and the fuel ratio (= fixed carbon/volatile matter weight ratio) FR [–] as:

a0 = (– 43FR + 164)N (ppm)

[15.14]



a1 = (22.8FR – 25.0)N (ppm/%)

[15.15]



a2 = (– 43FR + 219)N (ppm)

[15.16]

a3 = (0.81FR – 1.59)N (ppm/K).

[15.17]

and

The accuracy of this empirical correlation was ±50 ppm for eight types of coal or coal mixtures (Sakuno et al., 2002). Equation [15.13] includes only three operating parameters: oxygen concentration, sulphur capture, and bed temperature. Other parameters such as load, total pressure, and flue gas temperature were excluded because only three independent parameters were available during operation of 71 MWe PFBC. As shown in Fig. 15.2, the oxygen concentration, total pressure, flue gas temperature, and bed height are dependent on the plant load. The empirical equations presented above are qualitatively consistent with experimental findings reported by others, increase in NOx with increasing partial pressure of O2 (Andersson et al., 1999) and decrease in NOx with elevating pressure (Wallman et al., 1993; Koskinen et al., 1995). For the relation among N2O, NO, CO, and SO2 emissions, Abe et al. (2000b) proposed an empirical equation (ASH(TR)-SN equation):

(PN2O/PSO2)(PCO/PNO)2 = exp(43.7 ¥ 103/TC – 50.4).

[15.18]

The emissions of NOx from 71 MWe PFBC ranged from 70 to 350 ppm for coal types with 1.2–2.1 daf% nitrogen (Sakuno et al., 2002). This emission level is too high for operation without additional NOx abatement technology. Selective catalytic NOx reduction (SCR) reactors in the back pass after the gas turbine have been used (Fig. 15.1). In addition, selective non-catalytic reduction (SNCR), i.e., NOx reduction through gas-phase reduction by injecting NH3 into hot flue gas, has been conducted (Sakuno and Ueda, 2000; Shimizu and Ito, 2001). The SNCR reaction is an equimolar reaction of NO and NH3 as

4NO + 4NH3 + O2 Æ 4N2 + 6H2O.

[15.19]

It is interesting that NH3 injection was effective for NOx reduction, even at a low flue gas temperature of 650°C at 50% load. A NOx reduction of 25% was achieved by ammonia injection at a molar ratio of NH3/NOx =

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0.5. In contrast, SNCR using NH3 under atmospheric pressure is known to be effective only at temperatures higher than 800–850°C. This high SNCR efficiency observed for PFBC is attributable to increased reaction rates of homogeneous reactions under elevated pressure conditions. The combination of SNCR and SCR for deNOx was put into practice in Osaki 250 MWe PFBC (Shimizu and Ito, 2001). The NOx concentration of about 182 ppm at the furnace exit was reduced by 35% by SNCR at NH3/NOx = 0.5 and by 85% by SCR at NH3/NOx = 0.95 to meet levels required by regulation: 19 ppm (Shimizu and Ito, 2001). An advantage of PFBC is low N2O emissions at a high thermal load. Emissions of N2O are strongly affected by the freeboard temperature, as indicated by Eqs [15.10] and [15.12]. Such strong influence is partly explained by the homogeneous decomposition of N2O in the freeboard. Johnsson et al. (1992) determined the rate of decomposition of N2O in inert gas (N2) as follows:

d[N2O]/dq = –kN2O [N2O]CM 14 

kN2O = 6.2 ¥ 10 exp (–28230/T) (mol

[15.20] –1

3

cm s)

[15.21]

where T, q, and CM respectively represent temperature (K), gas residence time (s), and concentration of the third body M (inert gas such as N2) in gas phase (mol cm–3). Assuming a plug flow reactor, the N2O concentration at the outlet given by the rate constant, residence time, and the concentration of the third body is:

[N2O] = [N2O]0 exp (–kN2OCMq),

[15.22]

where [N2O]0 is the concentration at the inlet. The rate expression above indicates that the N2O decomposition is accelerated by elevated pressure because of increased CM. The concentration of N2O at the outlet of freeboard of a PFBC is given as Eq. [15.22] and N2O concentration at the bed surface if N2O formation in the freeboard is negligible. As shown in Fig. 15.11, the effect of temperature on N2O emissions from different PFBCs can be well explained by the homogeneous decomposition kinetics irrespective of plant size and operating conditions. These results suggest that N2O emission from PFBC can be well described by emissions from the bed followed by thermal decomposition as

[N2O] = [N2O]bed exp(–kN2OCMq),

[15.23]

where emission from the bed is given as a function of independent operating parameters such as

[N2O]bed = F3([O2], hS, Tbed).

[15.24]

However, such analysis has not yet been done. It remains as a subject of future work. As shown in Fig. 15.11(b), a discrepancy of the trend was observed

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Pressurized fluidized bed combustion (PFBC)

50

Calc: ([N2O] = [N2O]0 exp(–KN2OCMq) [N2O]0 = 60 ppm, Gas res. time q = 2 s Exper. Calc.

40 N2O (ppm)

685

P = 0.8 MPa

30 20 10 0 1050

60

1100 1150 Temperature (K) (a)

Calc: ([N2O] = [N2O]0 exp(–KN2OCMq) Gas res. time q = 4 s

N2O (ppm)

50 Calc., [N2O]0 30 = 70 ppm 40

20 10

1200

Exper. Black lignite Exper., HVB1 Exper., HVB2 LVB P = 1 MPa

Calc., [N2O]0 = 60 ppm

hN2O (–)

0 1020 1040 1060 1080 1100 1120 1140 Temperature (K) (b) hN2O = hN2O,0exp(–kN2OCMq), hN2O,0 = 0.07 Gas res. time q = 4 s 0.12 Exper. 0.1 Calc. ASH (TR) 0.08 0.06 0.04 0.02 0 900

50% load, 0.61 MPa O2 6.68%

75% load* 0.8 MPa, O2 4% 100% load 1.02 MPa O2 2.65%

1000 1100 Cyclone inlet temperature (K) (c)

1200

15.11 Effect of temperature on N2O emission or conversion of fuel-N to N2O (hN2O) from PFBCs of different size: (a) experimental data by Suzuki (2000); (b) experimental data by Andersson et al. (1999); (c) experimental data by Tsuji et al. (1999); solid lines – N2O emission/ conversion calculated using Eqs [15.21] and [15.23] assuming constant N2O emission/conversion from dense bed and thermal decomposition in gas phase in freeboard; broken line – ASH(TR) empirical equation (Eq. [15.10]); * pressure and oxygen concentration assumed by the present author.

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for low-volatile bituminous (LVB) fuel. For low-volatile fuels (high fixed carbon fuels), combustion of elutriated char fine particles in the freeboard cannot be neglected, as discussed in Section 15.4.2. A model incorporating combustion in the freeboard is therefore necessary for such fuels.

15.3

Development of combustion processes and technology

15.3.1 Coal-fired power generation Coal-fired PFBC power plants have been constructed and operated as demonstration projects and as commercial plants for two decades. Operating conditions of some PFBC units are presented in Table 15.1. The gas turbine output was 15–30% of the total output. This minor role of the gas turbine is attributable to the low flue-gas temperature. For PFBCs, the gas temperature is only about 1,120 K, but natural-gas fired gas turbines can achieve about 1,750 K. Most of the power from PFBC is produced through steam turbines. The steam condition is of great importance in achieving high efficiency. Most PFBCs have used subcritical steam conditions, but the largest PFBC, Table 15.1 Operating conditions of some PFBC demonstration and commercial plants Place

Tidd

Vartan

Wakamatsu Tomatouatsuma

Ohsaki

Karita

Country

USA

Sweden

Japan

Japan

Japan

Japan

Demo./Com.a

Demo.

Com.

Demo.

Com.

Com.

Com.

Total output [MW]

70.3

145

71

85

250

360

  ST output [MW] 55.5

110

56.2

73.9

213.5

290

  GT output [MW] 14.8

34

14.8

11.1

36.5

70

Feed system

Paste

Paste

Paste

Dry

Paste

Paste

Steam pressure [MPa]

9.0

13.7

10.2

16.7

16.6

24.1

SH temp [°C]

496

530

593

566

566

566

Bed height [m]



3.8

3.5



4

3.3–3.5

Pressure [MPa]

1.07

1.2b

1.1

0.95c

0.959c

1.3

Reference

Mudd and Almqvist Reinhart et al. (1995) (1991); Dahl (1993)

Misawa et al. (1999)

Kaneko et al. (1999)

Shimizu Koike and Ito et al. (2001) (2003)

a b c

Demonstration plant/Commercial plant. Pressure in pressure vessel. Pressure at GT inlet.

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Karita 360 MWe, adopted a supercritical steam condition. Karita PFBC is also the first supercritical steam fluidized bed boiler, including atmospheric FBCs. Using supercritical steam, its net efficiency becomes as high as 41.8% (HHV basis). This value is higher than that of pulverized coal combustion under similar steam conditions, e.g., efficiency of 40.3% with a steam condition of 24.13 MPa, 593°C (Koike et al., 2003). The combination of a high steam condition with a gas turbine makes PFBC a high-efficiency combustion technology.

15.3.2 Wet sludge combustion for compressed air production Pressurized wet sewage sludge incineration has been developed as a novel application of pressurized fluidized bed combustion in Japan by Tsukishima Kikai Co., Ltd, Sanki Engineering Co., Ltd, the Public Works Research Institute, and the National Institute of Advanced Industrial Science and Technology (AIST) (Nagasawa et al., 2008; Murakami et al., 2009). The concept is presented schematically in Fig. 15.12. Combustion of dewatered sewage sludge takes place in a pressurized fluidized bed reactor at elevated pressures up to 0.2 MPa (gauge) at a bed temperature of 1,000–1,100 K. The hot flue gas from the combustor is cooled to about 800 K by a gas–gas heat exchanger to preheat combustion air. Then the flue gas is filtered to remove the dust before it enters a gas turbine. The gas turbine drives an air compressor to produce compressed air, a part of which is used for combustion. One feature of this process is that the volumetric gas flow rate increases with combustion of wet sludge because of water evaporation and high temperatures. Consequently, the excess power is useful to produce Reactor

Heat exchanger

Filter

To flue gas treatment

Surplus air

Sludge

Bed Compressor

Gas turbine

Air

15.12 Pressurized sludge incineration process to recover compressed air.

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surplus compressed air, which is useful for wastewater aeration processes: the energy consumption for wastewater aeration treatment can be reduced. Another advantage of this concept is the low operation pressure. The reactor can be constructed without an expensive pressure vessel. Furthermore, this process can be cost-effective using a turbocharger for diesel engines (e.g., for automobiles) instead of a gas turbine–air compressor system. Turbochargers are readily available. They are not costly because of mass production. Even if some trouble occurs with the turbocharger, it can be replaced easily, quickly, and cheaply. Using a turbocharger, the electricity consumption to drive fans can also be reduced considerably (by about 40%). The emissions of N2O from pressurized combustors were reported to be lower than those from conventional sludge incinerators. The reduced N2O emission was attributed to the formation of a high-temperature region in the freeboard (for details, see Section 15.2.2). A three-year pilot plant test project subsidized by New Energy and Industrial Technology Development Organization (NEDO), Japan, was conducted from 2005. A pilot plant with inner diameter of 0.70 m, outer diameter of 1.20 m, and height of 9.2 m was constructed and operated. The incineration rate was 4.3 t dewatered sludge per day. The project finished successfully. Recently, Tokyo Metropolitan Government chose to introduce this technology for use in a waste water treatment plant.

15.4

Advantages and limitations of PFBC

15.4.1 Efficiency and plant cost As discussed in Section 15.3.1, PFBC can provide higher power generation efficiency than that of pulverized coal firing, given the same steam conditions. One disadvantage of PFBC, however, is its necessity for a heavy (and therefore expensive) pressure vessel. For example, the diameter, height, and the thickness of the pressure vessel of Karita 360 MWe PFBC are 15.4 m, 44.6 m, and maximum 0.15 m, respectively. The weight of the pressure vessel with the boiler is reported as nearly 3,600 tons, of which the weight of the boiler itself is only 800 tons (Koike et al., 2003). Because of its complicated structure, the cost of the PFBC boiler becomes high. Projected cost estimations by the Research Institute of Innovative Technology for the Earth, Japan (RITE, n.d.) set the cost per unit of power generation of PFBC as ¥320, 000/kW, whereas the cost of pulverized coal firing boiler with ultrasupercritical steam with net HHV efficiency of 40.9% is only ¥230,000/ kW. The cost of the PFBC plant is even higher than that of integrated coal gasification with combined cycle power generation (IGCC) of net efficiency of 43.0% (1300°C class gas turbine) to 47.0% (1500°C class gas turbine), each of which costs about ¥270 000/kW. Therefore, further cutting of the plant cost is necessary for PFBC technology. © Woodhead Publishing Limited, 2013

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The efficiency is limited by the gas turbine inlet temperature and the steam conditions, the latter of which is a common point with other combustionbased technologies. The gas-turbine inlet temperature is usually lower than that in the bed and the highest bed temperature is limited so that ash should not melt and required desulphurization efficiency can be achieved. Consequently, regarding the temperature given by the combustion condition, a further increase in gas turbine output is regarded as difficult to achieve. One approach to improve the efficiency is to add a gasifier (or pyrolyzer) to produce combustible gas from the same fuel and to burn the combustible gas in the flue gas from PFBC at the inlet of the gas turbine to increase the temperature. Sulphur is captured in the gasifier in the form of CaS. Then it is transported to the combustor with unreacted char to be oxidized to CaSO4. However, a detailed account of this technology is beyond the scope of this chapter. This subject will not be pursued further here.

15.4.2 Fuel flexibility Fuel flexibility is widely acknowledged as an advantage of FBCs. Not only PFBCs but also AFBCs can burn various fuels, such as anthracite, which cannot be used in pulverized coal firing. However, for some fuels, special care is needed for PFBCs. An important limitation is the fuel feed to the reactor operated at elevated pressures. A fuel–water mixture (paste) feeding system or a rock-hopper system for dry feed is necessary. Consequently, fuels must be pre-treated to meet the requirements (size and shape) of these feeding devices. The type of fuel affects combustion efficiency. To predict the combustion efficiency of fuel, hC (%), an empirical correlation was proposed by NBC (IEA Grimethorpe) (Clark et al., 1989) as:

1.7l 3 100 – hC = {48100/(Vqbed l )}{(1298 – Tbed)/Tbed}2.

[15.25]

where, V, qbed, l, and Tbed respectively represent the volatile matter content of the fuel (daf%), gas residence time in the bed (s), the stoichiometric air ratio (–), and the bed temperature (K). It is noteworthy that this equation does not include the total pressure term, possibly because the decreased diffusivity of gas with increasing pressure might cancel out the increase in oxygen partial pressure. The experimentally obtained results obtained by two pilot-scale plants of 71 MWe (Fig. 15.13) show similar trends to those suggested by Eq. [15.25], although Eq. [15.25] somewhat underestimates the combustion efficiency. The discrepancy is regarded as partly attributable to combustion of fine char particles in the freeboard. Carbon burn-up in the freeboard is affected both by chemical kinetics and mass transfer from the gas to the particles. For very fine particles of less than about 50 mm diameter, the reaction rate is controlled mainly by chemical

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Combustion efficiency (%)

100 99 98 Tbed = 1143 K Tbed = 1128 K Tbed = 1113 K Andersson et al. (1999) Tomuro et al. (1995)

97 96 95

0

20

40 60 VM content (daf%) (a)

80

Combustion efficiency (%)

100 99 98 97 Tbed = 1133 K

96 95

Wakamatsu 71 MWe 0

20

40 60 VM content (daf%) (b)

80

15.13 Effect of fuel type (volatile matter content) on combustion efficiency: comparison between (a) pilot-scale results for gas residence time of 3–3.4 s (Andersson et al., 1999; Tomuro et al., 1995) and (b) half-load condition of 71 MWe PFBC (Goto, 1998) with empirical correlation (Eq. [15.25]) (calculation condition: l = 1.2 and qbed = 3.4 s for pilot scale and l = 1.5 and qbed = 2.1 s for 71 MWe PFBC).

kinetics. The temperature of the fine particle is nearly the same as that of the surrounding gas because of its large specific surface area (external surface area/volume). For coarser particles, the reaction rate is often governed by oxygen diffusion through the gas film. Because the superficial velocity in the freeboard is about 1 m/s, the carry-over size of particles is estimated as about 0.25 mm. Consequently, kinetics-controlling and diffusion-controlling chemical conditions should be discussed. Assuming the operation parameters (pressure, freeboard temperature, and O2 concentration) used for the Wakamatsu 71 MWe shown in Figs 15.2 and 15.3, burn-up of fine carbon in the freeboard with gas residence time of 4 s is estimated. For chemical kinetic controlling conditions, the specific reaction rate of carbon, rC (kg-C m–2 s–1), was reported by Smith (1978) as

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Pressurized fluidized bed combustion (PFBC)



r = 3050 exp (–Ea/RT)PO2m,

691

[15.26]

where Ea, R, and m respectively represent the activation energy (=179.4 kJ mol–1), gas constant (J mol–1 K–1), and reaction order (0.5 or 1). For carbon with internal surface area of 100 m2 g–1, complete combustion of carbon is expected when the freeboard temperature becomes higher than 1,000 K, as shown in Fig. 15.14(a), but less reactive char with less internal surface

Mass transfer, 0.15 mm Mass transfer, 0.20 mm Kinetic, 100 m2/g Kinetic, 50 m2/g

0.8 0.6 0.4 0.2

Freeboard temperature (K)

0 40

60

80

Load (%) (a)

100

1150

2.5

1100

2

1050

1.5

1000

1

950

Temp.

Diffusivity of O2 (10–5 m2/s)

Carbon burn-up in freeboard (–)

1

0.5

Diffusivity of O2 900 40

60

Load (%) (b)

80

0 100

15.14 (a) Estimated burn-up of carbon carried over from bed to freeboard and (b) parameters for the estimation (char particle density of 1,000 kg-C/m3, reaction order m of 0.5, and Sherwood number of 2 are assumed).

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area (50 m2) necessitates a higher temperature for complete combustion. For coarser particles, for which the reaction is likely to be controlled by diffusion, particle size plays an important role: for smaller particles (0.15 mm) the burn-up is greater than 75%, but for coarser particles (0.2 mm) the burn-up was 41% at best. It is noteworthy that the diffusivity of oxygen in the freeboard gas is estimated as decreasing concomitantly with increasing plant load because the influence of pressure on diffusivity (inversely proportional to total pressure) overcomes the influence of increasing freeboard temperature (proportional to T1.5) under PFBC operation conditions. The combustion of fine particles in the freeboard plays an important role in determining not only the carbon burn-up efficiency but also the flue gas temperature. Yoshioka and Takezaki (2000) reported that at a half load (bed height = 1.8 m) about 8–10% combustion of the fed fuel occurred in the freeboard. They also reported that combustion in the freeboard decreased concomitantly with increasing of the load to less than 1.5% at full load (bed height = 4 m). Such a high combustion ratio in the freeboard under a partial load condition is attributable to reduced carbon burn-up in the bed, as described by Eq. [15.25] and high combustion rate under high oxygen partial pressure conditions. A high combustion ratio in the freeboard was also found when high-carbon content fuels, such as anthracite, were used. Figure 15.15 shows the change in freeboard and cyclone temperature of Karita 360 MWe PFBC with increasing anthracite content mixed with bituminous coal (Okuhata et al., 2004). Such an increase in the flue gas temperature affects the output balance of the gas turbine/steam turbine. At a bed height of 3.2 m (partial load condition), the GT output during bituminous coal combustion was about 61 MWe, but it increased to 72 MW when 100% anthracite was used. Experience related to the Karita 360 MWe PFBC revealed that special care must be taken sometimes for stable operation depending on the fuel properties (Okuhata et al., 2004). Accumulation of hard and coarse (>3 mm) ash in the coal (called ‘Zuri’) in the bed inhibits smooth fluidization and heat dispersion. It eventually results in the formation of sintered grains. With the formation of sintered grains, the bed’s apparent density decreases, becoming lower than 790 kg/m3. To avoid this problem, it was necessary to monitor the bed density and keep the content of ‘Zuri’ in the bed material lower than 3%. The control of the content of ‘Zuri’ was conducted by withdrawing the bed material and adding fresh material. For one kind of coal (Blair Athol coal), addition of Mg(OH)2 was effective for preventing sintering problems. This measure is regarded as changing the fusion temperature of the mineral matter. Such treatment was not necessary for the same coal in the case of Wakamatsu 71 MWe PFBC. The different behavior of coal ash is attributable to the difference in pressures: 1.3 MPa for Karita 360 MWe, but 1.0 MPa for Wakamatsu 71 MWe. The increase in pressure is regarded as increasing the heat release rate, i.e., increased risk of hot-spot formation.

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Pressurized fluidized bed combustion (PFBC)

Temperature (K)

1150

693

Fuel: mixture of anthracite and bituminous

1100

1050

Prim. cyc. temp. Freeboard

1000

3.7 m 0

50 Anthracite content (%)

100

15.15 Rise in freeboard temperature with anthracite combustion observed in Karita 360 MWe PFBC (Okuhata et al., 2004).

15.4.3 Attrition As discussed in Section 15.2.1, a considerable amount of sorbent was lost by attrition. Attrition affects not only the sorbent utilization efficiency but also the bed material size distribution, the latter of which affects the heat transfer coefficient of the boiler tubes in the bed. The heat transfer coefficient is known to increase with decreasing particle size, but it reduces the gas turbine output. The bed height required for given heat recovery rate decreases. Therefore, boiler tubes above the bed surface cool the flue gas, thereby reducing the gas turbine inlet temperature. To predict the bed material size distribution, information related to the attrition rate dependence on particle size is necessary. Saastamoinen and Shimizu (2007) analyzed the particle size distribution of the bed material of a 71 MWe PFBC and proposed a rate expression of attrition as

ddp/dt = –kidpi,

[15.27]

where dp, i, and ki respectively represent the particle diameter, exponent, and attrition rate constant for exponent i. The respective values of i and k2 were 2 and 8.5 ¥ 10–4 m–1 s–1. It should be emphasized that this attrition rate expression (second-order dependence upon size) and rate constant are valid only for the limestone used for Wakamatsu PFBC. The feature of limestone used for Wakamatsu PFBC was that primary fragmentation after injection into a hot fluidized bed was negligible. Consequently, only attrition was regarded as occurring in this plant. However, it is very likely that primary fragmentation by thermal shock occurs depending on the limestone type. For such a case, a model of the particle size change should take account of fragmentation followed by attrition. Different values for exponent n (0–2.6) have been reported for limestone of different types (Franceschi

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et al., 1980). Fragmentation and the attrition rate as well as their dependence upon particle size must be evaluated experimentally for each type of limestone. The second-order dependence of limestone attrition on particle size suggests that there exists an optimal particle size of the fed limestone to suppress the attrition/carry-over loss. The attrition rate will decrease with decreasing feed size whereas contents of fine particles in the feed will also increase, which will result in the increased carry-over. Consequently, a limestone crusher that can produce particles with a narrower particle size distribution will be favorable. However, this is a matter related to crusher specifications. It will not be discussed further here.

15.4.4 In-situ SO2 capture Combustion of high-sulphur fuels without an external wet scrubbing process is a salient advantage of FBCs. Therefore, sulphur capture in PFBCs using high-sulphur coal is of great concern. The theoretical approach discussed in Section 15.2.2, however, suggests that the apparent reaction order of desulphurization by uncalcined limestone (CaCO3) with respect to SO2 concentration is less than unity when attrition occurs with sulphation. This low dependence implies that the sulphur capture rate becomes insufficient when high sulphur content fuel is used. The experience of Wakamatsu 71 MWe PFBC burning mixtures of petroleum coke (S content 3.5%) and coal revealed that sulphur capture was determined mainly by the feed rate of the surface area of fresh limestone larger than carry-over size, but not by the surface area of bed materials (Sakuno et al., 2001; Saastamoinen and Shimizu, 2007). Their results imply that the removal rate of the product layer at the surface of the bed material via attrition was insufficient for SO2 capture, and that the surface of fresh limestone reacted mainly with SO2. For such cases, one approach to improve the sulphur capture efficiency is the reduction of the limestone feed size, although elutriation of fine particles will occur. Another approach is to recycle some fly ash into the bed so that fine limestone particles (smaller than carry-over size) can have sufficient residence time in the freeboard. The results of fly ash recycling of Wakamatsu 71 MWe PFBC showed that the surface area of fresh fine limestone particles (75–250 mm) was nearly the same as the surface area of coarse particles (>250 mm). However, these measures will affect the operation of given PFBC. The reduced bed material size will increase the heat transfer rate of the boiler tubes in the dense bed. It will reduce the bed height to meet the steam generation requirement, thus it will reduce the gas temperature in the freeboard (Fig. 15.3). Finally, it will reduce the gas turbine output. Indeed, the bed height of the 71 MWe PFBC with fly ash recycling was about 10% lower than that without fly ash recycling at the same load (Fig. 15.2(a)).

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It should be emphasized that the change in one parameter might affect the overall operation of the PFBC.

15.4.5 Boiler tube damage and plant availability Any fluidized bed boiler that has in-bed tubes cannot be free from the risk of tube erosion. PFBC is no exception. Kyushu Electric Power Co., Inc., which has been operating Karita 360 MWe PFBC since 2001, has reported some erosion difficulties. One erosion-related difficulty of boiler tubes near the bottom occurred after about 25,000 hr of operation. The mechanism was reported to be the direct contact of boiler tubes with the upward flow of fuel particles. This kind of trouble is not regarded as specific to PFBCs. The countermeasure was hardening of the boiler surface (Kyushu Electric Power Co., Inc., 2006). Another problem is specific to PFBCs. A plug on the furnace wall dropped off and compressed air in the pressure vessel flowed into the bed because the pressure of the surrounding air in the vessel is higher than that of the bed. The solid flow induced by the air stream caused the erosion of boiler tubes. The countermeasure was to fix the plug by welding instead of screwing (Kyushu Electric Power Co., Inc., 2001a, 2001b). Although some difficulties have been reported as described above, Karita PFBC achieved 25,000 hr of operation in 2006 and 48,000 hr operation in 2011. During operation, a continuous operation record of 3,411 hr was achieved in 2003.

15.5

Conclusion

Pressurized fluidized bed coal combustion for power generation is currently the most efficient power generation technology based on direct combustion of coal. This technology can be especially advantageous for application in regions with low water availability because of its characteristic in-situ sulphur capture without a wet scrubbing system. Emissions of NOx can be suppressed with ammonia injection by combining SNCR at elevated pressures and SCR at ambient pressure. Emissions of N2O can be suppressed by thermal decomposition in the freeboard at elevated pressure. One problem for PFBCs is their high plant cost. Another problem is boiler tube erosion in the bed, although this problem is shared with other fluidized bed technologies. Further development of high-efficiency coal-fired PFBCs is needed. Aside from coal combustion, pressurized sewage sludge incineration with smaller reactors operated under moderate pressure conditions has been proposed. This sewage sludge PFBC is expected to reduce energy consumption of wastewater treatment and to reduce greenhouse gas (N2O) emissions. This sewage sludge PFBC is expected to be put into operation soon.

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15.6

Fluidized bed technologies for near-zero emission combustion

Sources of further information and advice

The books listed below are useful for understanding PFBC technology: Cuenca, M.A. and Anthony, E.J. (eds) Pressurized Fluidized Bed Combustion, London: Blackie Academic & Professional (1995). Rousaki, K. and Couch, G., Advanced Clean Coal Technologies and Low Value Coals, IEA Clean Coal Centre, CCC/39 (2000). Wu, Z., Understanding Fluidized Bed Combustion, London: IEA Clean Coal Centre, CCC/76 (2003). Wu, Z., Developments in Fluidized Bed Combustion Technology, London: IEA Clean Coal Centre, CCC/10 (2006).

15.7

References

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15.8 a i C C M d p D e E a FR i k i k N2O k S m M n N

Appendix: notation coefficients for Eq. [15.13] (sensitivity factor for operating parameter) concentration of SO2 [kmol m–3] concentration of the third body M in gas phase [mol cm–3] particle diameter [m] effective diffusivity of SO2 through product layer [m2 s–1] activation energy [kJ mol–1] fuel ratio (= fixed carbon/volatile matter weight ratio) [–] exponent of attrition rate expression (Eq. [15.27]) [–] attrition rate constant for exponent i [m1–i s–1] rate constant of homogeneous N2O decomposition in inert gas [mol–1 cm3 s] rate constant of sulphur capture by uncalcined limestone [kmol(1–n) m(2–3n)] reaction order of carbon combustion [–] molecular weight of CaCO3 [kg kmol–1] reaction order of SO2 capture by uncalcined limestone nitrogen content of coal [daf%]

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nitrous oxide concentration [mol cm–3] or [ppm] oxygen concentration in the flue gas at SCR inlet [vol.%] partial pressure of carbon monoxide [atm] partial pressure of nitrogen oxide [atm] partial pressure of nitrous oxide [atm] partial pressure of oxygen [atm] partial pressure of sulphur dioxide [atm] specific reaction rate of carbon combustion [kg-C m–2 s–1] specific reaction rate of sulphation by uncalcined limestone [kmol m–2 s–1] rsteady specific reaction rate of SO2 capture by uncalcined limestone under steady sulphation–attrition condition [kmol m–2 s–1] R gas constant [J mol–1 K–1], T temperature [K] Tbed bed temperature [K] T C flue gas temperature at the cyclone [K] V volatile matter content of fuel [daf%] [N2O] [O2] PCO PNO P N2O P O2 PSO2 r C r S

15.8.1 Greek symbols a d h C h S l q qbed r t

average attrition rate [m s–1] product (CaSO4) layer thickness at the surface of uncalcined limestone [m] combustion efficiency of fuel [%] sulphur retention efficiency [–] stoichiometric air ratio [–] gas residence time [s] gas residence time in bed [s] density of limestone [kg-CaCO3 m–3] interval of intermittent attrition [s]

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