Circulating fluidized bed combustion (CFBC)

Circulating fluidized bed combustion (CFBC)

16 Circulating fluidized bed combustion (CFBC) W. N o w a k and P. M i r e k, Czestochowa University of Technology, Poland DOI: 10.1533/978085709880...

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16

Circulating fluidized bed combustion (CFBC)

W. N o w a k and P. M i r e k, Czestochowa University of Technology, Poland DOI: 10.1533/9780857098801.3.701 Abstract: This chapter is devoted to basic aspects of circulating fluidized bed combustion (CFBC). The chapter first reviews the history and current status of CFBC technology. The chapter then discusses advantages and basic principles of CFBC, describing factors affecting combustion efficiency as well as the issue of reliability and availability of CFB boilers. The chapter includes a development strategy and challenges of CFBC technology. Key words: circulating fluidized bed (CFB) boiler, circulating fluidized bed combustion (CFBC), pollutants, fuel.

16.1

Introduction

16.1.1 History and current status of circulating fluidized bed combustion (CFBC) technology In the large number of books devoted to scientific progress over the years, one major conclusion can be drawn: the most important inventions and discoveries of our times have been made by accident. This is also true for the circulating fluidized bed (CFB). In 1938 Warren Lewis and Edwin Gililand of MIT (the Massachusetts Institute of Technology) developed a new gas-solid process, when they were trying to find a proper gas–solid contacting process for fluid catalytic cracking (FCC). They applied the term ‘fast fluidization’ as another mode of fluidization and referred to it as the ‘upflow operation’ to distinguish this process from bubbling fluidization. The concept of a coalfired fluidized bed boiler was first proposed by Odel from the US Bureau of Mines in 1943, after his invention of FCC, referred to above. Early in the 1950s Houdry and Combustion Engineering began testing Odel’s concept for a coal-fired fluidized bed boiler. Unfortunately, these tests were halted due to solids handling problems. The idea of burning fuel in a fluidized bed was again recognized by Douglas Elliott of CEGB, UK in the early 1960s. In 1965, a US engineering company, Pope, Evans and Robbins, under a contract with British Coal Utilisation Research Association (BCURA) and the Central Electricity Generation Board (CEGB), commenced combustion tests with a 12≤ ¥ 16≤ circulating fluidized bed test rig. The development and improvement of fluidized bed technology have been 701 © Woodhead Publishing Limited, 2013

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made by a number of groups worked independently, including among others: Lurgi, Foster Wheeler and Ahlstrom Pyropower. The first Lurgi CFB boiler designed for the production of electricity and heat was built in 1982 at the Vereingte Aluminium Werke at Luenen, Germany. This plant generated 84 MW total (9 MW electricity, 31 MW process steam, 44 MW molten salt melt), burning low-grade coal washery residues with added limestone. In 1990 Lurgi put into operation a 150 MWe CFB boiler for the Texas New Mexico Power Corporation in the US, with a water-cooled cyclone separator. A water-cooled cyclone was also used on the 240 MWth Bewag AG plant in Berlin, Germany (Hilger, 1991). Foster Wheeler pioneered the use of fluidized bed technology in the 1940s but in the early 1970s launched a research and development program into CFB technology at its John Blizard Research Center in Livingston, New Jersey. This resulted in the design basis for the world’s first commercial fluidized bed steam generator (4 t/h steam), the CFB boiler supplied to Georgetown University in 1979. The company’s experience with large-scale CFB boilers began in 1987 with the 110 MWe Tri-State Nucla power project of Colorado-Ute Electric Association (Fig. 16.1) (Nucla, 1995, 1999). Foster Wheeler’s CFB experience has matured in projects in the US, Europe and the Far East, in applications that include coal combustion of low grade fuels, resulting in many CFB boilers in operation and under construction (Fig. 16.2). These include the largest CFB boiler currently in operation (the Lagisza 460 MWe supercritical CFB boiler located in Poland), which currently is the only operational supercritical CFB boiler in the world. Foster Wheeler’s latest pioneering and most advanced supercritical CFB boilers with total steam power 4 ¥ 550 MWe will be seen in Samcheok, South Korea in 2015. The Ahlstom Pyropower circulating fluidized bed (CFB) was developed in the late 1970s. Since 1969, CFB research had been one of the major research and development projects of the Hans Ahlstrom Laboratory in Karhula, Finland. Further investigation of CFB technology was initiated in the mid-1970s. A pilot plant was constructed in Karhula, Finland in 1976. The first commercial CFB boiler by Ahlstrom Pyroflow commenced operation in January 1979 at Pihlava, Finland. The second Ahlstrom Pyroflow CFB was the Kauttua co-generation CFB plant commissioned in May 1981. Other CFB boiler manufacturers active in the market included: Stein Industrie, which had a cooperation agreement with Lurgi, Tampella Power, Kvaerner Pulping Power Division and Babcock & Wilcox (B&W). Tampella Power (a Finnish Company) is best known for its 160 MW CFB boiler CYMIC, with a cyclone located inside the combustion chamber delivered in 1994 to United Paper Mills’ Rauma Mill in Finland. Before 1994 Kvaerner Pulping Power Division supplied Bodens, Torvvarme with a 20 MW CFB using a water-cooled cyclone in 1985 and in 1986–88 supplied ADM, USA

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Saturated steam to superheaters CFB boiler

Steam drum

Main steam

Secondary superheater

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Hot cyclone

Original turbines (3 ¥ 12 MWe)

New turbine (74 MWe) Generator

Final superheater Primary superheater Economizer Boiler feedwater return

Hot cyclone

Generator

Cooling water Cooling water

Deaerator

Waterwall tubing

Deaerator

Fly ash Air preheater

Secondary air Primary air

Loop seal Dual combustion chambers

Coal Limestone

Baghouse Induced draft fan

Fly ash

Fly ash Bottom ash

16.1 Schematic diagram of NUCLA CFB system (adopted from Nucla, 1995).

Stack

704

Fluidized bed technologies for near-zero emission combustion 3000

Capacity (MWe)

2500

CFB boilers delivered over the period 2006–2012 Coal

CFB boilers under construction over the period 2010–2015

2000

1500 Petcoke 1000

500

0

Bituminous Sub-bituminous coal

Texas lignite

Renewable fuels

CFB boiler’s primary fuel

16.2 Total capacity of Foster Wheeler’s CFB boilers fired with different types of fuels delivered over the 2006–2012 period and under construction over the period 2010–2015.

and Cedar Rapids, IA, with hot (i.e. not water- or steam-cooled) cyclones generating 8 ¥ 54–60 kg/s steam. Babcock & Wilcox had concentrated its development of CFB boilers in the area of so-called impact separators (U-beams located immediately downstream of the furnace). The B&W CFB boilers uses this separator system without cyclones. ABB Combustion Engineering utilized technology from Lurgi GmbH. The examples of ABBs Combustion Engineering projects by 1994 were, among others, Scott Paper Co., 1986 (294 t/h), New Brunswick Electric, 1986 (95 t/h); Texas-New Mexico, 1990 (2 ¥ 499 t/h, 447 t/h); Northeastern Pwr PA, 1989 (220 t/h); Ultrasystems, 1989 (2 ¥ 132 t/h); Schuykill Energy PA, 1989 (374 t/h); AES Thames, 1989 (2 ¥ 304 t/h); Colmac Mecca, 1992 (2 ¥ 105 t/h) (Skowyra and Tanca, 1993). Although each company refined its own CFB process, there are two main technical options with regard to the CFB process: ∑ ∑

Lurgi technology (also used by ABB Combustion Engineering); and Foster Wheeler (FW) technology.

Lurgi CFB boilers used an external heat exchanger (FBHE), whereas Foster Wheeler CFB (before 1994) did not. The differences between these two systems can be seen in Fig. 16.3. Lurgi CFB boilers use a CFB loop with an external heat exchanger (FBHE) where about 20–60% of the total heat (excluding the boiler back pass) is absorbed. As far as hydrodynamics of the process is concerned, both processes (FW vs Lurgi) are similar. The FW

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process used a simple circuit without any external heat exchanger. Most of the heat is absorbed in the bed. Tubes located in the FBHE absorb heat from the circulating solids particles. Solids flowing through the FBHE can be easily controlled and as a result the bed temperature can be maintained at a given level. In recent years, a large numbers of different designs of CFB boilers have been available in the market. In general, they are classified on the basis of arrangement and types of cyclone separators and circulation of water–steam mixture into first and second generation of CFB boiler design. The practical example of a first generation CFB boiler design is two 300 MWe CFB boilers at Jacksonville Energy Authority in Jacksonville, Florida, USA. These units are the largest CFB boilers based on natural circulation. The second generation design of CFB boilers was introduced in 1992 by introduction of water/steam-cooled integrated COMPACT separators, and in 1996 by use of INTREXTM heat exchanger located in the furnace. Recently, once-through supercritical (OTSC) CFB technology has been developed, enabling the next

Cyclone

Superheater

Economizer Feed water

Air-heater

Furnace

Secondary air Coal and limestone

Baghouse

L valve FBHE

Stack Fly ash

Bottom ash

Primary air (a)

16.3 CFB Technology: (a) Lurgi with external FBHE, (b) Foster Wheeler with internal heat exchangers.

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Live steam Superheater

Cyclone

Superheater

Economizer Air-heater Furnace Baghouse Secondary air Coal and limestone

Stack Fly ash Primary air Bottom ash (b)

16.3 Continued

stage in CFB development (Hotta, 2009). The advantages of supercritical CFB boilers in comparison to subcritical boilers are: ∑

lower capital costs – SC CFB boilers consume less steel due to their use of smaller diameter tubes and the absence of large steam drum (Basu, 2006), ∑ lower pressure drop over the furnace tubing, resulting in less power needed for feed water pumps and lower auxiliary power consumption (Hotta, 2009), ∑ homogeneous combustion temperature in CFB (both vertically and horizontally) (Hotta, 2009). The ongoing challenge for CFB technology is developing capacities as well as plant efficiencies as a response to an ever-growing requirement for power generation. As follows from Hotta (2009), in 2012 the technology is ready for scale-up of critical CFB components, such as the furnace, solids separators and fluidized bed exchangers to 800 MWe.

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16.1.2 Advantages of CFBC The advantages of CFBC technology can be considered from two points of view: process and environment. From the process point of view, the advantages of CFB technology can be summarized as follows: ∑ ∑ ∑ ∑ ∑ ∑ ∑

possibility of burning a diverse range in fuel properties, possibility of multi-fuel firing, stable operating conditions and good turndown ratio (Hotta, 2009), no need for fuel preparation (e.g. pulverizing) (Hotta, 2009), support firing is not needed except during start-up periods (Hotta, 2009), increased capacity possible within the same footprint as old boilers (Hotta, 2009), high bed-to-surface heat transfer coefficients, 100–400 W/m2K (Hotta, 2009).

As CFB technology becomes more widely accepted, and unit sizes increase, even more focus is being placed on environmental performance of this technology. In fact, the low stack emissions are a driving force that dictates the selection of CFB boilers over pulverized coal (PC) boilers.

16.1.3 Environmental advantages of CFB boilers The main environmental advantage of the CFB technology is its ability to burn a diverse range of difficult low grade fuels of varying quality with low emissions of NOx, low-cost sulfur capture during combustion in the furnace itself, as well as low CO and CxHy emissions due to turbulent conditions and good mixing. Additionally, CFB boilers allow the application of relatively simple solutions for the problem of NOx emission to the atmosphere. A key feature of CFB boilers is that the combustion temperature is low (generally 850°C) and air is introduced gradually, such that NOx emissions are low. Further parameters allowing for lowering the levels of formation and emission of NOx include: ∑ ∑ ∑ ∑ ∑ ∑

excess of air (oxygen concentration in the combustion chamber), relation between primary and secondary air, changes in boiler load, kind of fuel inserted into the boiler, pressure profile in the combustion chamber, and gas–solids mixing in the combustion chamber.

In CFB boilers in Poland, a decrease in boiler load leads to a decrease in temperature in the combustion chamber. This has an impact on NOx emission. For instance, in a 235 MWe CFB boiler, a 50% decrease in boiler load results

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in the upper part temperature dropping from 850°C to 700°C. NOx emission as a function of boiler load is shown in Fig. 16.4. The observed trends in emission levels are close to those in PC boilers equipped with low-NOx burners and air staging. When the load is decreased, the total excess air in the boiler must be increased. NOx emissions rise when the load falls to a certain level. This is a typical feature of low-NOx firing systems. However, the CFB combustor is capable of keeping NOx emissions below 300 mg/ Nm3. The unburned combustibles in the bottom and fly ash are generally kept to below 2%. CFB utility units have lower NOx emissions than grate and PC boilers because of the low uniform combustion temperature. Present low-NOx technology (low NOx burners, air staging, reburning) for PC boilers can only achieve emission levels that are twice those of a CFB boiler. On the other hand, the emission of nitrous oxide (N2O) from fluidized bed boilers is much higher (100–200 ppm) compared to that from PC boilers, where the emission of this gas is considerably lower due to the latter’s high combustion temperature (1200–1500°C) (Basu, 2006). In comparison to bubbling fluidized bed (BFB) boilers, CFB boilers give lower NOx emissions burning similar fuels because of inherent differences in the contacting pattern (Brereton, 1996). CFB boilers offer the possibility of removing the sulfur dioxide during the combustion process by the addition of limestone (CaCO3) directly to the combustion chamber. Using such limestone directly with the fuel inserted into the furnace leads to a high efficiency of desulfurization and eliminates the need 300

NOx emission (mg/Nm3)

280 260 240 220 200 180 160

100

120

140

160 180 Boiler load (MWe)

200

220

240

16.4 NOx emissions from 235 MWe CFB boiler fired with brown coal.

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for installing different kinds of scrubbing equipment. The limestone should not only have the right chemical composition, its particle size distribution should also be close to the appropriate particle size distribution for the boiler (see Fig. 16.5). However, achieving sorbent of a strictly defined particle size distribution is an extremely difficult task. Whilst the improvement of the particle size distribution of limestone significantly improves the desulfurization efficiency, the excess of limestone needed to ensure the desired level of desulfurization is frequently too high. In order to achieve 90% reduction of SO2 emission, the sorbent has to be added in the 2.5–4 range of Ca/S molar ratio, according to the type of fuel and sorbent (reactivity level, crystallinity, particle size distribution, hardness, etc.). It is desirable to achieve an optimally low level of Ca/S molar ratio to better manage the processes taking place within the fluidized bed. This also has an impact on limestone costs and the amount of ash subject to utilization. Figures 16.6 and 16.7 compare the desulfurization efficiency with Ca/S ratio for bituminous and brown coals. The data are presented for commercial CFB boilers in operation in Poland. It follows from Fig. 16.7 that bituminous coal requires less limestone to achieve the required desulfurization efficiency. However, in case of brown coal, a high level of desulfurization efficiency 0.99 0.98 0.95

Particle size distribution (–)

0.90 0.80 0.70 0.60 0.50 0.40 0.30 0.20 Recommended range 0.10 0.05 CFB 235 MWe Turow power plant Recommended

0.02 0.01 10

100

d (mm)

1000

10000

16.5 Recommended and real particle size distribution of limestone for 235 MWe CFB boiler operated in Turow power plant.

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SO2 desulfurization efficiency (%)

100

80

60

40 Sulfur content in brown coal: <= 0.35 >0.35 <= 0.55 > 0.55

20

0

0

2

4

6 8 10 Ca/S molar ratio (–)

12

14

16

16.6 Desulfurization efficiency in CFB boilers with brown coals.

SO2 desulfurization efficiency (%)

100

80

60

40 Sulfur content in bituminous coal: 20

<= 0.35 > 0.35 <= 0.55 > 0.55

0 0.0

0.5

1.0

1.5

2.0 2.5 3.0 3.5 Ca/S molar ratio (–)

4.0

4.5

5.0

16.7 Desulfurization efficiency in CFB boilers with bituminous coals.

requires an excessive amount of limestone and consequently a large amount of residues is produced. The search for more effective emission reduction techniques leads directly to the practical application of fluidized bed combustion for coal-fired plants in Poland. The guaranteed emissions from Polish commercial CFB

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plants are much lower even than from other CFB boilers, as can be seen in Table 16.1. All CFB boilers under operation in Poland must meet national environmental regulations and must be able to comply with the even more stringent regulations expected in the future. The CFB technology is a good solution for future environmental challenges like CO2 reduction both in repowering of coal-fired power plants and greenfield power plants. This technology is well suited for co-firing coal with biomass and burning a broad spectrum of CO2 neutral fuels such as waste-derived biomass fuels, proving its ability as an efficient and economic way of reducing CO2 emissions.

Table 16.1 Guaranteed emissions from selected CFB boilers Plant name

Start-up Fuel

Capacity

Guaranteed emissions mg/Nm3 (6% O2) NOx SO2

CO

Dust

Polpharma, Starogard Gdan´ski · Zeran´, Warsaw

1996

Bituminous coal 60.2 MWth

300

400

250

50

1997

Bituminous coal 315 MWth

200

200

250

50

Bielsko-Biała

1998

Bituminous coal 177 MWth

250

300

250

50

Gardanne, Provence

1996

Bituminous coal 250 MWe

250

400





Ceder Bay, Jacksonville, USA

1994

250* 350 Bituminous coal 3 boilers with total capacity 250 MWe

250



Tonghae, Korea

1998

Bituminous coals

200 MWe

515

515





Turów, Poland

1998

Brown coal

235 MWe

256

238



50

Jaworzno II, Poland S´wiecie

1999

Low-rank coals

Two 70 MWe 200* 200

150

50

2004

Coal, biomass

234 t/h coal 180 t/h biomass

200* 250 300 250

100 100

30 30

Elcho Compact

2003

Coal

2×104 MWe

246

383

200

50

Turów CFB Compact

1999– 2004

Brown coal

3×260 MWe

371

347

150

50

Lagisza SC-CFB

2009

Coal slurry

460 MWe

200

200

200

30

Fortum Czestochowa

2010

Coal, biomass

66 MWe

200

200

200

30

* ammonia injection.

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Generation and destruction of pollutants A key benefit of the CFB technology is low emissions allowing the strictest environmental standards to be met. Unfortunately, even the most efficient combustion process generates pollutants, which in the case of a fossil fuel-fired CFB boiler are: sulfur dioxide (SO2), nitric oxide (No), nitrogen dioxide (No2), carbon monoxide (Co), hydrocarbons and dust. Because the major product of coal combustion is nitric oxide, nitrogen oxide emission (Nox) is usually related to this pollutant. As can be seen from Fig. 16.2, the most popular are coal-fired CFB boilers. Therefore, the description of the mechanism of generation and destruction of pollutants will be limited to coal-fired CFB boilers. The formation of sulfur dioxide during the combustion of coal is the effect of sulfur oxidation according to: S + O2 fi SO2 + 296 kJ/g mol

[16.1]

a part of the sulfur dioxide may be absorbed by the calcium oxide contained in the mineral matter of coal forming the calcium sulfate according to: SO2 + 1 o2 + CaO fi CaSO4 + 486 kJ/g mol 2

[16.2]

another part of sulfur dioxide may be converted into sulfur trioxide according to: SO2 + 1 o2 fi SO3 2

[16.3]

The rest of the unconverted sulfur dioxide is typically discharged into the atmosphere, but in the case of CFB combustion, SO2 is captured with the help of a low-cost sorbent like limestone (CaCo3) or dolomite (CaCo3◊MgCo3). The retention of sulfur dioxide in the fluidized bed with limestone takes place in two steps: 1. endothermic reaction of calcination: CaCo3  CaO + CO2 – 183 kJ/g mol

[16.4]

2. exothermic reaction of sulfation [16.2] (Fig. 16.8). A key role in the calcination reaction [16.4] is played by the equilibrium partial pressure of carbon dioxide, at the calcination temperature (Stantan, 1983). The sulfation rate and sulfation degree of a limestone particle are the function of many parameters including: limestone particle size, structure, porosity after calcination, temperature, gas velocity around the particle and Cao content and impurities. The exact mechanism of sulfation reaction is still debatable and is a matter of ongoing investigations. The additional

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Circulating fluidized bed combustion (CFBC) CO2

CaO

713 CaO

486 kJ/g mol

–183 kJ/g mol Calcination

Sulfation SO2

CaCO3

CO2

CO2

CO2

CO2

CaSO4

16.8 Schematic illustration of calcination and sulfation reactions.

information on that subject can be found in anthony and Preto (1995), Dennis and Hayhurst (1990), oka (2004) and Johnsson et al. (1990). From the operational point of view, it is very important to determine the level of sorbent utilization. For a given level of sulfur capture, this parameter is a function of cyclone efficiency, reactivity of sorbent, its size and temperature and it can be estimated based on the following simplified expression (Basu, 2006): PU U (ASH ASH ) Ec ln (1 – Essoor ) Fsor 3.12 Esorrbav H f S – 100 P = Fc Ec [d m max ax XCaCO3 rbav H f + 100 PU ln(1 – Essor )]

[16.5]

where Fsor is the sorbent feed rate, kg/sec; Fc is the coal feed rate, kg/sec; Esor is the sulfur capture efficiency; rbav is the average bed density, kg/m3; S is the sulfur fraction in coal; Hf is the height of the furnace above the secondary air level, m; P is a proportionality constant in pore plugging time s (= P × (SO2 concentration, kmol/m3)–1); U is superficial gas velocity, m/s; Ec is average cyclone efficiency; dmax is maximum extent of sulfation; and XCaCO3 is the weight fraction of calcium carbonate in the sorbent particle. The sulfur capture efficiency in CFB boilers is influenced by a number of design and operating parameters, including: ∑ ∑ ∑ ∑ ∑ ∑ ∑ ∑ ∑ ∑

combustion temperature (oka, 2004), Ca/S ratio (Oka, 2004), residence time of limestone particles in the fluidized bed (Oka, 2004), primary to secondary air ratio (oka, 2004), excess air (oka, 2004), fluidization velocity (Oka, 2004), coal feed (oka, 2004), fly ash recirculation (Oka, 2004), characteristics of coal and ash (oka, 2004), combustor pressure (Basu, 2006).

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The second group of pollutants emitted in significant quantities during the combustion process from CFB boilers are nitrogen oxides. This term is related to so-called NOx emissions including in most countries two species: nitric oxide (NO) and nitrogen dioxide (NO2), and additionally nitrous oxide (N2O). Amongst these, nitrous oxide is recognized as a greenhouse gas which may contribute significantly to the greenhouse effect. From the fluid bed combustion point of view it is very important because emissions of nitrous oxide are much higher for low temperature combustion sources, such as CFB boilers, than from higher temperature sources such as PC boilers. Hence, it is expected that the level of N2O emission will be regulated in the future. The concentrations of nitrogen oxide depend strongly on: the coal type and its characteristics, excess air, bed and freeboard temperatures, the mass stream of secondary air, char hold-up in the bed and the type of inert material (Oka, 2004). Hence, understanding the mechanisms involved in the formation and destruction of nitrogen oxides is a key element in the search for the most efficient method to reduce these pollutants. Nitrogen oxides can be formed from nitrogen in the air or nitrogen in the fuel in three ways: ∑ ∑ ∑

in an oxidizing reaction between nitrogen and oxygen in the combustion air (giving ‘thermal’ NOx), in an oxidizing reaction of fuel-bound nitrogen (giving ‘fuel’ NO x), in reaction between molecular nitrogen and hydrocarbon radicals (giving ‘fast’ or ‘prompt’ NOx).

At typical fluidized bed combustion temperatures (800–900°C), thermal nitrogen oxides are negligible. They are significant for PC boilers, where temperature in the freeboard can achieve ca. 1,400°C. Hence, in CFB boilers most nitrogen oxides are the product of oxidation of nitrogen present in the volatile matter and char of fuel. As follows from the experimental results given in Suzuki et al. (1991), depending on coal type, 17–30% of the fuel nitrogen contributes to the formation of nitrogen oxide and nitrous oxide. During the formation and destruction of nitrogen oxides, a large number of complex chemical reactions are involved. There are over 80 possible reactions for NOx and N2O formation and over 90 possible reactions with HCN playing a significant role in nitrogen oxide and nitrous oxide formation from nitrogen released during devolatilization (Oka, 2004; Johnsson et al., 1990). Some of the reactions are catalyzed by calcined and spent limestone as well as char. In general, there are two separate paths for NOx and N2O formation: 1. Homogeneous oxidation reactions and gaseous compound reduction involving amines (NCO), hydrogen cyanide (HCN) and ammonium (NH3), released during coal devolatilization and pyrolysis, i.e.:

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HCN + O ´ NCO + H

[16.6]



NCO + O ´ NO + CO

[16.7]



NCO + NO ´ N2O + CO

[16.8]



NCO + OH ´ NO + CO + H

[16.9]



NH3 + 5/4 O2 = NO + 3/2H2O

[16.10]

2. Heterogeneous oxidation of nitrogen in char particles,

N – from char + O2 ´ N2O

[16.11]

or nitrogen reduction on the surface of char particle: N – from char + NO ´ N2O

[16.12]

Coals with higher volatile matter content produce considerably lower N2O concentrations and higher NOx concentrations and vice versa. The total amount of NOx and N2O is typically higher for coals with high volatile matter content. In the process of destruction of nitrogen oxides and nitrous oxide and the formation of nitrogen, bed solids have a significant influence, especially char. Destruction of N2O is possible by the following reactions:

N2O + H ´ N2 + OH

[16.13]



N2O + CO ´ N2 + CO2

[16.14]



N2O + OH ´ N2 + HO2

[16.15]

It can be concluded, that formation and destruction of nitrogen oxides and nitrous oxide are complex processes that can proceed: ∑ ∑

directly – from nitrogen in char, indirectly – by homogeneous reactions of gases from char and nitrogen compounds in volatiles (HCN, NH2, NH3) as well as heterogeneous reactions of gases and char (Oka, 2004).

The emission of nitrogen oxides in CFB boilers is influenced by a number of design and operating parameters, including: ∑ combustion temperature (Brereton, 1996), ∑ Ca/S ratio (Brereton, 1996), ∑ excess air (Brereton, 1996; Oka, 2004), ∑ primary to secondary air ratio (Oka, 2004). In investigation presented in Mirek et al. (2007), it was observed that the air nozzle design influences the temperature distribution in the combustion chamber. Experimental tests were carried out on two 670 MW circulating fluidized bed boilers. The basic constructional difference between these boilers

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was the use of different primary air nozzles. In the case of boiler A, the primary air grid was equipped with so-called ‘pigtail’ nozzles, while in the case of boiler B so-called ‘arrowhead’ nozzles were used. As shown in Fig. 16.9, the change in primary air distribution in the lower part of the combustion chamber has substantially influenced the intensity of the combustion process and the temperature distribution along the CFB structure height. A direct effect of the thermal and flow condition differences between both boilers tested were different CO, SO2 and NOx concentrations. The operating conditions of the boilers and the test results obtained are summarized in Table 16.2. As shown by Table 16.2, the basic difference occurred in the case of NOx emission. The concentration of this gas component amounted to 152 mg/mn3 for boiler A and 238 mg/mn3 for boiler B. It can be stated that for the two boiler units, supplied with the identical fuel (in terms of quality) 1180 Maximum design bed temperature, T = 1172 K

Temperature, T (K)

1160

1140

670 MW – CFB boiler A 670 MW – CFB boiler B

1120

1100

1080

Minimum design bed temperature, T = 1089 K

0

5

10 15 20 25 30 Distance from the gird, z (m)

35

40

16.9 Distribution of temperature along the combustion chamber height of 670 MW CFB boilers (Mirek et al., 2007). Table 16.2 Results of tests on the 670 MW CFB boilers (Mirek et al., 2007) Parameter

Boiler A 670 MW

Boiler B 670 MW

Q, W SO2, mg/mn3 (6% O2) NOx, mg/mn3 (6% O2) CO, mg/mn3 (6% O2) Ca/S, mol Ca/mol S l, – Dpf, Pa

627 ¥ 106 345 152 8 2.6 1.2 7,490

627 ¥ 106 363 238 3 2.6 1.2 7,520

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and running under comparable operating conditions, the factor determining the difference in the amount of emitted nitrogen oxides was the design of the air nozzles. High combustion efficiency in CFB boilers guarantees a low emission of carbon monoxide. Emission of this gas is a function of fuel characteristics and fuel distribution in the combustion chamber, that is thermal and flow conditions inside the furnace. In Polish CFB boilers, depending on fuel fired in the boiler, guaranteed emissions vary from 100 to 200 mg/Nm3 (6% O2) (Table 16.1). Additional information on the mechanisms of generation and destruction of sulfur dioxide and nitrogen oxide is given in Chapter 9. Utilization and disposal of solid wastes Combustion of solid fuels is always associated with the production of ash. Due to significant addition of sorbent for dry flue gas (FG) desulfurization, the relative mass of ash from FBC is larger than for pulverized coal combustors. In each case it is necessary to achieve an optimum level of Ca/S molar ratio (usually as low as possible); otherwise problems with bed hydrodynamics, increase of sorbent costs and the mass of ash which is then subject to utilization may become significant. It is clear that larger amounts of ash bring about significant problems for the power station, e.g. associated with storage costs, needs to search for and prepare large areas for storage, countermeasures that have to be taken to decrease the dust emission (mainly from open storage facilities) and the costs of maintaining the storage areas and its facilities (Kobyłecki et al., 2004). In the case of CFB boilers, mainly two types of ash, i.e. bottom ash (BA) and fly ash (FA), are produced. Both differ with the particle size distribution, content of unburned carbon and chemical composition. Their common feature is, however, that they contain the components present in the fluidized bed, i.e. fuel ash, unreacted and reacted sorbent used for desulfurization and unreacted coal/char. BA of particle sizes of roughly 0.3–6 mm (approximately 50% of solids have diameter <1 mm) usually contributes to 30–60% of the total ash produced. FA, captured by ESP and having particle size distributions of roughly 1–300 mm, contributes to 40–70% of the total ash mass. Apart from BA and FA, FBC also produce ‘ultrafine’ ash of particle size less than 5 mm, that are emitted to the atmosphere (usually at a concentration of roughly 10 ppm). The properties of fluidized bed ash depend on several parameters. The most significant are: fuel and sorbent characteristics, Ca/S ratio and combustion conditions as well as boiler type (sorbent residence time in the oxidizing zone, instantaneous concentration of SO2, combustion temperature, OFA inlet ports, air ration, boiler size, etc.). Accordingly, the properties of any

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particular ash have to be established case by case for a particular installation using particular fuel and sorbent. Example compositions of selected fluidized bed (FB) ashes are shown in Table 16.3. According to the ash properties proposed at the conference organized by the University of Kentucky in 1997, FB and PC ashes can be classified into three classes taking into consideration their pozzolanic and hydraulic properties: 1. F – containing <8% of CaOfree and having pozzolanic properties, 2. CI – containing 8–20% of CaOfree and having both pozzolanic and hydraulic properties, 3. CH – containing >20% of CaOfree and having hydraulic properties. Good pozzolanic properties, and the presence of CaOfree and CaSO4 in FBC ash may result in proper control of the hydration and production of stable and strong mineralogical structures. However, due to its exothermic reaction with water, which brings about material expansion, FB ash does not meet the conditions defined by ITB #328, PEN-450, EN-450, ASTM C-618-92a and ASTM C-618-89a standards in order to be used for cement or concrete production. Accordingly, it has to be treated individually case by case in order to propose and accept a technology for its application/reuse, e.g. in civil engineering. Fluidized bed ashes seem to be an interesting resource for various industries. Taking into consideration a strategy of FB ash utilization, the following options are being considered: ∑

Reuse of raw (untreated) ash. This does not require any investment costs and the ash can be used for neutralization of heavy metals in sewage sludge (after grinding to a surface area of roughly 4,000 cm 2/g), agglomeration of coal enrichment products for FBC, to stabilize soil, for agriculture industry (recultivation, de-acidation, bacteria-killing agents, soil structure modifiers), in civil engineering (as low-class concretes), in mining and drilling industry, in geotechnology (fillers), in hydro-industry, as whitening agents, for macronivellations (e.g. as for Makro Cash & Carry, Poland) or in synthetic industry. ∑ Reuse of modified/treated ash (e.g. by selective separation of some components, hydration, etc.). Here the ash can be used for cement production, mixed with PC ash or slags, in concrete industry or as soil stabilizer or in the brick industry. According to EU standard EN-450, in order to be used in concretes, the FA has to have the following properties: LOI < 5%, SO3 < 3% (<5% according to ASTM C–618); CaOfree < 2.5%, and particle size distribution <34% on 45 mm sieve. Since usually SO3 > 3% in FB ashes, the criterion can be fulfilled by mixing the FA with PC ash.

© Woodhead Publishing Limited, 2013

Table 16.3 Composition of fluidized bed ashes (Kobyłecki et al., 2004) © Woodhead Publishing Limited, 2013

LOI (wt%)

SiO2 (wt%)

Fe2O3 (wt%)

Al2O3 (wt%)

CaOfree in MgO CaOtotal (wt%) (wt%)

SO3 (wt%)

Na2O (wt%)

K 2O (wt%)

PFBC, Canada, mixture of BA & FA

14.24

11.16

8.42

43.53

17.32

0.34

15.75





CFB, France, BA CFB, France, FA

12.00 24.46

40.60 13.47

2.77 11.21

4.90 6.20

29.17 29.87

9.7 8.82

0.05 0.52

10.33 14.05

– –

– –

CFB, Canada, BA CFB, Canada, FA

5.08 24.021

19.87 12.00

4.28 11,96

5.75 6.00

44.23 30.66

15.34 9.93

0.59 0.42

17.71 14.11

– –

– –

CFB, Germany, BA CFB, Germany, FA

12.74 1.07

26.74 28.72

6.81 17.92

17.00 12.54

19.42 21.20

4.43 7.06

2.75 2.73

10.13 12.14

2.08 1.12

1.68 1.01

CFB, Poland, FA CFB, Poland, BA-fine CFB, Poland, BA-coarse (>6mm)

10.70 2.20 2.60

38.90 57.40 61.10

5.70 3.0 3.80

1870 1680 25.50

13.70 10.10 0.80

2.90 2.80 0.12

2.30 1.50 1.60

6.00 6.10 0.50

0.68 0.43 0.50

1.41 1.53 2.45

0.94

33.45

5.62

27.34

18.26



1.61

7.66

0.68

2.11

CFB, Poland, BA

5,51

CaOtotal (wt%)

720



Fluidized bed technologies for near-zero emission combustion

Reuse/application of specially treated ash. The product can be used for special binders, mineral substitutes in the concrete industry for foamconcrete production. An interesting option is also the use of ‘specially treated’ FB ash, Flubet®, for grout curtains in river embankments (applied to embankments in some areas along the Olza and Odra rivers in Poland), curtains protecting groundwater against pollution (material UTEX-GMM), diaphragm walls, or at hydro engineering embankments. This ash-cement mixture is: relatively cheap, as only 5% of cement is added; self-adjusting to the local soil conditions; characterized by a large angle of internal friction (associated with good mechanical and static properties of the embankment); has low water penetration; and low heavy metals leaching to the groundwater, which makes it a good material to stop leaching from refuse dumps, etc.

Three main technological approaches are being considered in order to apply FB ash in the cement or concrete industry: ∑

Controlled hydration (CaO becomes Ca(OH)2) + preparation of ash–other components mixture in order to optimize the pozzolanic properties of the product. ∑ Mixing + grinding of FBC FA, slag and Portland clinker in order to capture anhydrite and prevent the formation of ettringite in the concrete. The process is not affected by the ash composition, but requires strict control of the amount of CaO and CaSO4 in the ash. Unfortunately, the quality of the product is pretty poor. ∑ Processing taking into consideration hydraulic and pozzolanic properties of the FA – based on the fact that the additives for homogenization and concrete production are chosen. The process is promising for ashes of 5–10% CaOfree and the product is of high quality. As briefly described above, a relatively wide range of possible profitable applications of FB ash is possible. However, based on the majority of research conducted in Poland, four main ways to utilize/reuse FB ash seem to be the most promising. These are application of the ashes: ∑ as materials for road construction (Glinicki, 2001), ∑ in the cement industry as a cement substitute, ∑ to produce zeolites for capture of CO2, SO2 or other components (Majchrzak-Kuceba, 2001), ∑ to produce new sorbents for flue gas desulfurization (Szymanek, 2002). The main benefits of such applications of FBC ashes are lower concrete manufacturing costs, lower CO2 emission, less sorbent used, lower operational costs for the power station, as well as various benefits for the environment, associated with the production of cleaner energy. © Woodhead Publishing Limited, 2013

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16.1.4 Fuel flexibility One of the main advantages of the CFB technology is the ability to burn a wide range of fuels. The CFB has the flexibility to handle fuels with varying sulfur content, whilst meeting the required emissions levels. The ability of CFBs to burn a wide variety of fuels, whilst meeting strict emission control regulations (see Table 16.1), makes them an ideal choice for burning such fuels as high sulfur coal, lignite, peat, oil, sludge, petroleum coke, gas and wastes. All of these fuels are burned cleanly and economically in CFB boilers, without the need for complex scrubbers, catalytic or other costly chemical clean-up systems. The Polish experience with CFBs covers the firing of bituminous and subbituminous coals, brown coal and coal slurry. Typical bituminous coals are easy to burn in a CFB boiler due to their consistent heating value, moderate volatile content and high ash softening temperature. Most of the CFB units in Poland have fired bituminous coals and brown coal. The average lower heating value of brown coals and coal slurry ranged from 8 to 12 MJ/kg. The moisture content in the case of coal slurry was up to 37%. CFB units in Poland have fired bituminous coals ranging up to 2.4% sulfur (CFB boiler OFz-425 Siersza) and brown coal with sulfur content up to 1%. High ash fuels are suitable for CFB applications, with no adverse effect. Coal slurries with ash content of up to 53% (two 70 MWe Compact CFB boilers at Jaworzno II and 120 MWe at Katowice) have been burned together with bituminous coal, resulting in a good performance. In addition, not-so-typical fuels such as bark and paper sludge have been utilized in a bubbling type fluidized bed boiler (EC Ostrołęka). Biomass with a heating value of 6.78 MJ/kg and moisture content up to 65% was burned effectively in a CFB-234. The MCR of this unit when burning coal only was 234 t/h (efficiency 92%) and when burning biomass 180 t/h (efficiency 90.8%). Low volatility fuels can also burnt in CFB boilers. The hot circulating particles provide a large ignition source for low volatile fuels, resulting in excellent flame stability and burning characteristics. CFB boilers can also effectively treat coal washery waste and other refuse that are generally not utilized in PC boilers. The price of these low quality fuels is very low, making CFB technology very competitive with other combustion techniques.

16.1.5 Comparison of CFBC technology and alternative combustors Circulating fluidized bed combustion is nowadays counted as one of those technologies which facilitates the use of fuels in an environmentally satisfactory

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and economically viable way. Thanks to that, CFBC technology belongs to the group of clean coal technologies (CCTs) together with: ∑

pulverized coal combustion (PCC) with supercritical steam driving a steam turbine, together with flue gas cleaning units; ∑ pressurized fluidized bed combustion (PFBC), which currently uses bubbling bed boilers, in combined cycle with both a gas and steam turbine; ∑ integrated gasification combined cycle (IGCC), which uses different types of gasifier, and in combined cycle with both a gas and steam turbine; ∑ combined heat and power (CHP) applications, where the (subcritical) steam turbine is designed to produce both power and useful heat for process or district heating. One of the great features of pulverized coal-fired boilers is that they can operate under supercritical (SCPC), advanced supercritical (ASCPC) and ultra-supercritical cycles (USCPC). Conventional PC boilers operate at much higher temperatures (approximately 1350–1650°C) compared to the CFB boilers. The CFBC at the lower temperature has several benefits: ∑ ∑

lower sorbent requirement, fuel ash never reaches its softening or melting points (900ºC is considerably below the ash fusion temperatures of most fuels). Hence, the fouling and slagging problems that are characteristic of PC units are significantly reduced, if not eliminated. Moreover, the design of a CFB boiler is not as dependent on ash properties as is a conventional PC boiler, because combustion temperatures are below ash fusion temperatures, ∑ lower temperature reduces nitric oxide emissions by nearly eliminating thermal NOx.

Commercial CFB units offer greater fuel diversity than PC units. Moreover, the use of larger fuel sizing in CFB boilers reduces auxiliary power and pulverizer maintenance requirements and eliminates the high cost of pulverizer installation. On the other hand, the power requirements of the primary air fans for the CFB boiler are higher than that of the primary air fans for a PC boiler. Hence, one can say that the auxiliary power requirements for PC and CFB technologies are relatively similar. Overall efficiency of PC boilers is higher (currently 47% at Power Plant Nordjylland 3, Denmark, with a target of 50–55%) than that of CFBC. Moreover, the emission of nitrous oxide (greenhouse gas) from PC boilers is significantly lower in comparison to CFB boilers. Currently, the capacity of PC-fired boilers exceeds a limit of 1000 MWe, while in the case of CFB boilers the largest unit operating under supercritical cycle in Lagisza Power Plant has a capacity of 460 MWe. Pressurized fluidized bed combustion is based on known components working in a combined advanced clean-coal system: fluidized bed, gas

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turbine and steam turbine. Taking into consideration P200 combustor, the combustion of coal takes place in the bed vessel at about 850°C under high pressure (12 bar). PFBC technology is a one-step process for converting coal directly into power, with a number of advantages over conventional CFBC technology such as: ∑ ∑ ∑ ∑ ∑ ∑ ∑

higher total plant efficiency resulting from the combined-cycle power generation, higher combustion efficiency resulting from the burning of fuel under high pressure (12–16 bar), compact design, utilization of low-grade fuels, utilization of wet fuels, smaller amount of solid and gaseous wastes per unit of electricity produced, useful ash products.

In comparison to conventional atmospheric combustion plant, the flue gas from PFBC combustors has lower mass flow, higher total pressure, higher CO2 pressure and very low flue gas oxygen content. This makes PFBC technology well suited for CO2 capture in combination with other technologies, e.g. the Benfield process. In investigation presented in Bryngelsson and Westermark (2009), it was observed that in the Sargas process for carbon dioxide capture, the CO2 removal efficiency is greater than 98% with CO2 quality greater than 94%. The capacity of PFBC boilers is lower than that of CFBC boilers. The world’s largest PFBC boiler with capacity 360 MWe is currently operated in the Karita Plant in Japan and uses ABB Carbon P800 technology. Integrated gasification combined technology (IGCC) is a combination of two technologies: coal gasification, which uses coal to create syngas (cleanburning gas) and combined-cycle, which is the most efficient commercially available method of producing electricity. IGCC has the potential to use coal in a more efficient process and with lower emissions than PC and CFB boilers, obtaining a value of 50%. Moreover, IGCC technology results in significantly lower water consumption and solids production as well as the production of a large number of by-products for industrial use like hydrogen, sulfur and methanol. The real advantage for IGCC comes in relation to carbon capture and sequestration (CCS). Unfortunately, the technical and economic performance of IGCC depends more strongly on feedstock quality than in CFB systems. Moreover, IGCC plants cannot operate with such a wide variety of fuels as in CFB boilers. Another drawback of IGCC technology in relation to CFB technology is reliability. In the case of IGCC plant, the annual availability is in the range of 80–85% on coal versus approximately 90–92% for coal-fired CFB boiler. Although IGCC is an advanced technology that represents the cleanest of currently available coal technologies, the power

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generation cost as well as still low availibility of IGCC plants are the main barriers to implement this technology on a larger scale.

16.2

Basic principles of circulating fluidized bed combustion (CFBC)

The process of coal combustion under CFB conditions is significantly different from the processes which take place in other types of boilers. Complete understanding of this process requires a knowledge of many complex processes connected with chemistry, hydrodynamics and heat and mass transfer in fluidized beds as well as characteristics of a fuel injected into the combustion chamber of a CFB boiler. Since the most popular are the coal-fired CFB boilers, the description of the combustion process will focus on this type of fuel. A fresh coal particle introduced into a combustion chamber filled with hot inert material particles passes through various stages of physical, physicochemical, and chemical transformations, which typically occur simultaneously. After entering the fluidized bed, the coal particle temperature rises rapidly and when a temperature of 100°C is attained, the first process of intense evaporation and drying begins. after drying and heating of coal particles, the next step during the combustion process is pyrolysis or devolatilization. This process is caused by thermal degradation of hydrocarbons and depends upon factors such as type of coal and its particle size, initial and final temperature, pressure, exposure time at the final temperature as well as heating rate. The process of volatile matter releasing begins at temperatures around 500°C and can be influenced by the furnace hydrodynamics. In fact, it consists of two overlapping processes: volatile releasing and volatile combustion. The volatiles released from an isolated particle can burn around it but when a large number of particles are injected into the bed, the volatiles from adjacent particles join together, forming plume (Stubington and Clough, 1997). The rate of the devolatilization process can be estimated assuming that thermal decomposition of coal particles can be described as the sum of a great number of simultaneous independent chemical reactions. The quantity of a particular component released during pyrolysis may be represented by the first order reaction (Oka, 2004): dV i dVM = kRi (VM VM i* – V VM Mi) dt

[16.16]

where: VM is amount of volatile yet to be released, %; VM* is total volatiles content, %; i is independent reaction for ith component of volatiles, which has specific activation energy, i.e., reaction rate; and t is time, s. The rate of reaction kRi can be expressed in the form of the arrhenius law:

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Circulating fluidized bed combustion (CFBC)

Ê Ei ˆ kRi = kRoi Roi exp Á – Ë RgT ˜¯

725

[16.17]

where: kRo is pre-exponential coefficient, s–1; E is activation energy, kJ/mol; Rg is universal gas constant, 8.314 kPa m3/kmol k, and T is temperature, k. The amount of volatiles left in char, at the moment t, is obtained by integrating expression [16.17]: Ê VM i* – VM i = VM i*exp Á Ë

Ú

t

0

ˆ kRi ddt˜ ¯

[16.18]

Introducing the assumptions that, for all reactions the same coefficient kRoi may be used, and that the activation energy E may be represented by the continual distribution function f (E), the total quantity of volatiles left in char can be calculated by integrating the expression [16.18] in the whole interval of activation energies:

Ú

VM * – VM = VM *



0

È exp Í– Î

Ú



0

˘ k (E ) dt ˙ f (E ) dE ˚

[16.19]

By introducing the assumptions that f (E) is a Gaussian distribution and the integration limit is of –• to +•, the quantity of volatiles released at a time t is given by (anthony et al., 1974): VM * – VM =

¥

Ú

+• –•

VM * s G (2p )1/2

È exp Í– kRo Î

Ú

t

0

Ê ˆ (E – E0 )2 ˘ exp Á – E ˜ dt – ˙ dE Ë RgT ¯ 2s G2 ˚

[16.20]

where sG is the standard deviation of the Gaussian distribution of activation energy, kJ/mol. evaporation and formation of volatiles inside the particle cause fuel particle disintegration (or fragmentation). This situation takes place when the pores in the coal particle are not capable of conducting all of the evolved gaseous matter to the particle surface and the inner pressure breaks the particle into several pieces (oka, 2004). From that moment on, the char (or devolatilized fuel) combustion process begins. Combustion of char takes place in the pores of the particle or at its surface and starts after or during the evolution of volatiles from the parent coal particle. in general, the combustion of char is a two-step process which involves transportation of oxygen to the carbon surface and reaction of carbon with oxygen on the carbon surface (Basu, 2006). The total combustion rate

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of char particles in the fluidized bed is influenced by the character and the ratio of the rates of consecutive processes such as (Oka, 2004): ∑ ∑ ∑

oxygen diffusion from the bubbles to the emulsion phase, oxygen diffusion in the emulsion phase to the char particle surface, oxygen diffusion through the particle’s porous structure to the carbon molecules at the surface of the pores, ∑ heterogeneous carbon oxidation occuring at the surface, ∑ diffusion of oxidation products occuring in the reverse direction to the bubbles with simultaneous homogeneous reaction in the gaseous phase. These rates depend on the combustor’s operating condition and characteristics of the char particle. There are a large number of different mathematical models describing the combustion process of char particles (Avedesian and Davidson, 1973; Basu et al., 1975; Campbell and Davidson, 1975; Horio and Wen, 1978; Chen and Saxena, 1978; Ross and Davidson, 1981; Turnbull et al., 1984). They differ by applying different combustion regimes, oxidation reactions, fluidized bed models and mass-transfer models. The additional information on the mechanisms of generation and destruction of sulfur dioxide and nitrogen oxide is given in Chapter 9.

16.3

Circulating fluidized bed (CFB) boiler process and performance

In commercial CFB boilers, achieving full capacity is the most important requirement. The requirements concerning emission levels are also important if the boiler must operate to strict environment standards. In some CFB boilers there were some difficulties to operate the boiler properly. The main problem was that the boiler could not always receive a full load or sometimes the temperature was too high. This situation was also associated with a large temperature difference between the upper and lower part of the combustion chamber. Generally, the high temperature in the combustion chamber was due to poor particle size distribution (PSD) of the bed material and low separation efficiency. The boiler capacity depends on the heat exchange rate in the CFB unit. In a CFB loop the heat is mainly transferred by solids particles. The total value of heat, which must be exchanged in the combustion chamber can be even 70% of the heat created during the combustion process. This value depends on the boiler construction and fuel type and is also a function of the fuel’s LHV. The main way to change the value of heat flow is to change the solids mass flow carried up in the combustion chamber. This can be controlled, by changing the solids mass flow reintroduced into the combustion chamber through the return leg. In addition, the high efficiency of desulfurization requires sufficient solids circulation. To achieve a low NOx

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727

emission, the temperature in the upper and lower parts of the combustion chamber must be almost uniform, which again is achieved by a proper solids circulation. The above issue is connected with the PSD in the fluidized bed which, together with the solids mass flow and the fuel’s LHV, is the one of the crucial parameters affecting the performance of the CFB boiler.

16.3.1 Operating conditions in the combustion chamber of a CFB boiler The relation between fuel solids mass flow, fuel LHV and temperature profile in the combustion chamber of the CFB boiler can be analyzed by using a simple thermal balance for a CFB as shown in Fig. 16.10 (Nowak and Bis, 1997). The CFB loop in Fig. 16.10 is divided into three blocks (I, II, III). Block No. i is referred to as the lower part of the combustion chamber, block No. ii is referred to as the upper part of the combustion chamber with heat exchangers, and block No. iii is referred to as the convective section. For each boiler section, a thermal balance equation can be written. The energy flow of solids in each section is defined as follows: upper 0 Qcir ∫ hcyc · m f · RR · H cir

[16.21]

where m f is mass flow of fuel introduced into the furnace, kg/s; and RR is recirculation ratio, QciI r ∫ m f · RR · H clower ir

[16.22]

Qcioutr ∫ (1 – hcyc ) · m f · RR · H cout ir

[16.23]

Q inf ∫ m f · LHV LH

[16.24]

Q If ∫ m f · LHV LHV · (1 – 0.6l I /l )

[16.25]

The mass flow of solids, which are carried up in the chamber can be expressed by using a recirculation ratio (RR) defined by: [16.26] RR ∫ m·cir/mf where m·cir is mass flow of circulating solids, kg/s. The ratio of solid flow enthalpy in the upper part and lower part of the r lower chamber H ciuppe /H cir can be obtained from the model as a function of solids r recirculation ratio RR with LHV as a parameter. The enthalpy ratio roughly r lower indicates the temperature profile in the chamber because H ciuppe /H cir is r upper lower proportional to the temperature ratio Tcir /Tcir . Figure 16.11 shows the

© Woodhead Publishing Limited, 2013

Fluidized bed technologies for near-zero emission combustion

728

Qfg upper Hcir = cmT upper

Qb

II

Qcir

hcyc

QE

III Luvo

Qa

II

0

Qa

out

Qfg

II

Qa

I

Qfg lower Hcir =

I

I

Qf Qcir

out

Qcir

cmT lower in

Qf

Qba

0

Qcir

I

I

Qa

Key: 0 Qcir – energy of circulating solids flowing into lower part of combustion chamber, kW



I

Qcir – energy of circulating solids flowing into combustion chamber, kW

II Qcir out Qcir I Qf in Qf

– energy of circulating solids flowing into a cyclone separator, kW – energy of circulating solids flowing out of the convection section, kW – energy of fuel flowing into lower part of combustion chamber, kW – energy of fuel flowing into combustion chamber, kW



Qfg – energy of flue gas flowing into the convective section, kW



Qfg – energy of flue gas flowing into combustion chamber, kW



I

out Qfg

– energy of flue gas flowing out of the convective section, kW



Qa – energy of air flowing out of the LUVO, kW



Qa – energy of air flowing into the LUVO, kW



Qa – energy of primary air flowing into lower part of combustion chamber, kW



0

I

II Qa



– energy of secondary air flowing into combustion chamber, kW

QE – energy of heat exchanger in combustion chamber, kW Qb – energy of heat exchanger in the convective section, kW Qba – energy of the bottom ash flowing out of the bottom part of the combustion chamber, kW

Hcir

upper

– solid flow enthalpy in the upper part of the combustion chamber, kJ/kg

lower Hcir

cm Tupper T lower hcyc

– solid flow enthalpy in the lower part of the combustion chamber, kJ/kg – specific heat capacity, J/kg K – temperature in the upper part of the combustion chamber, K – temperature in the upper part of the combustion chamber, K – cyclone efficiency, –.

16.10 A simplified thermal outline of a CFB boiler (Nowak and Bis, 1997).

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Circulating fluidized bed combustion (CFBC)

729

1.2 LHV = 18 MJ/kg

upper lower Hcir /H cir (–)

1.0

A B

0.8 LHV = 23 MJ/kg

0.6 0.4

LHV = 25 MJ/kg

0.2 0.0

1

tt = 850°C hcyk = 0.95 l1/l = 0.5 l = 1.2 thermal output = 100%

10 RR (–)

100

16.11 Enthalpy ratio of solids in the upper and lower part of the combustion chamber versus RR.

results calculated for CFB Żerań OFz-450. The values of LHV = 18 MJ/kg and LHV = 25 MJ/kg are an example of changes of fuel LHV. r lower The desired solution can be found at the intersection of the H ciuppe /H cir r upper lower = 1 vs. RR curve and the line H cir /H cir = 1. The interval between r lower points a and B on the H ciuppe / H = 1 line is the intersection area for cir r the possible range of LHV. It corresponds to the range of RR, which should be practically assured to keep a required uniform temperature profile. In the present example, the maximum calculated value of RR is about 50.

16.3.2 Critical importance of correct particle size distribution The PSD of solids has a fundamental influence on the recirculation ratio and accordingly on other output and emission parameters of a CFB boiler. Under commercial operations, it is difficult to keep a constant PSD in a boiler’s chamber mainly due to attrition, erosion and combustion processes in the fluidized bed (Nowak and Bis, 1997; Report EVT, 1997; Report PIE, 1997). The above difficulties were experienced in the CFB boilers like CFB Żerań in Poland. The comparison of solids PSD in the Żerań boiler during its initial period of operation is shown in Fig. 16.12. As can be seen in Fig. 16.12, the PSD curves of coal, sorbent and ash do not match up with the values requested by EVT/RAFAKO. The greatest disagreement can be seen for fluidized bed ash, which is the main make-up medium in the bed. The operational problems were analyzed and found to be connected with this disagreement (data for 1997/1998) (Nowak and Bis, 1997; Report PIE, 1997). Key to improving that situation was a change of the PSD in the fluidized bed into the required ones. The first idea was crushing

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Fluidized bed technologies for near-zero emission combustion

730

Cumulative weight fraction (–)

1.0

d50 of the cyclone

0.8

Crushed ash

Silica sand

Coal

0.6 Limestone Ash

0.4

0.2

0.0

Lim. required by EVT/RAFAKO 0.1

Coal required by RAFAKO

dp (mm)

1

10

16.12 PSD of solids in CFB boiler OFz-450 operated in Z˙eran´ Power Plant. Lines indicate PSD required by EVT/RAFAKO, and lines with points are the actual ones.

the bottom bed ash. Nevertheless, even such a modified ash did not have the required PSD. The next idea was a recirculation of fly ash (RFA) and fine particles from the bottom ash (RBA). Such changes improved boiler output, but were expensive and difficult to operate, and in fact, did not guarantee keeping the right emission levels all the time. in such situations, the solution to the problem can be the detailed analysis of the influence of PSD of inert material on the boiler performance, which can be carried out with the help of the population balance of solids in the bed. The population balance proposed takes into consideration combustion, hydrodynamics of the bed and mutual interactions between particles during fluidization. Three streams of solids in the fluidized bed were taken into consideration: ash, sorbent and coal. For each stream the steady state balance equations for the kth class of particles can be written as follows: ∑

for ash stream: ash m inash (dk ) – m eash (dk ) + m Rash (dk ) + m RFA (dk ) – m Dash (dk ) ash + m RD (dk ) – m aastth (dk ) + m iasn h (dk +1) = 0



[16.27]

for coal stream: c m inc (dk ) – m ec (dk ) + m Rc (dk ) + m RFA (dk ) – m Dc (dk ) c + m RD (dk ) – m actt (dk ) + m actt (dk +1) – m ccom (dk ) = 0

© Woodhead Publishing Limited, 2013

[16.28]

Circulating fluidized bed combustion (CFBC)



731

for sorbent stream: s m ins (dk ) – m es (dk ) + m Rs (dk ) + m RFA (dk ) – m Ds (dk ) s + m RD (dk ) – m astt (dk ) + m astt (dk +1) = 0

[16.29]

where m is mass flow rate of a specified solid fraction, kg/s; c is coal, s is sorbent, com is combustion, dk is diameter of the kth class of particles, e is elutriation, R is return through the return leg, RFA is recirculation in fly ash, D is drainage, RD is recirculation in drainage, and att is attrition. in the population balance model the following expressions were used: ∑

for coal fragmentation based on Bellgardt et al. (1987), the fragmentation coefficient of dk th class is equal to: n + 2ˆ D1/fr3 ÊÁ n1 2 Ë 2 ˜¯ j= Ddk



–1/3

dk

where Dfr is fragmentation factor, –; n1 is primary fragmentation coefficient, –; n2 is secondary fragmentation coefficient, –; and dk is diameter of the kth class of particles, m. Probability of fragmentation of the dkth class: Pfr (dk ) = 1 – e–(dk /d ffrr )/w ffrr



0.56Arg mbcomp

·

U g2 Ut (dk )

[16.32]

where A is boiler cross-section area, m2; rg is flue gas density, kg/m3; mbcomp is mass of component in the bed, kg; Ug is flue gas velocity, m/s; and Ut is terminal velocity, m/s. attrition rate based on arena et al. (1983): katt (dk ) = K att



[16.31]

where wfr is fragmentation coefficient, –; and dfr is specific diameter, m. on the basis of general expression to calculate the elutriation rate constant (Davidson et al., 1985), the following expression was used in the model: ke (dk ) =



[16.30]

Ug · dk 3

where Katt is attrition constant, 1/m. Char combustion rate based on Field et al. (1967):

© Woodhead Publishing Limited, 2013

[16.33]

732

Fluidized bed technologies for near-zero emission combustion

kcc (dk ) =



2ppg 1 · r c 1 (d ) + 1 kf k kc

where pg is oxygen partial pressure, Pa; rc is coal density, kg/m3; kf is mass transfer coefficient, s/m; and kc is combustion rate constant for char, s/m. Combustion rate constant for char: MR Rg Ts , kc = ko e– E /M ko = 0.888 s/m, E = 152,000 J/kkmol mol



[16.34]

[16.35]

where ko is pre-exponential factor for char combustion, s/m; E is activation energy, J/kmol; Ts is temperature of the coarse fuel particle, K; Rg is universal gas constant, J/mol K; and M is molar mass, g/mol. Oxygen mass transfer coefficient: jDdiff M dk ˆ 48jD Ê k f (dk ) = Á1 + ˜ 2d (dk )¯ dk RT RTgs Ë

[16.36]

where j is index of particle class after fragmentation, –; Ddiff is oxygen diffusion coefficient, m2/s; d is thickness of surface boundary layer, m; and Tgs is gas temperature in surface boundary layer, k. Numerical calculations were conducted concerning the influence of the following parameters: mean diameter ‘d50’ of coal and sorbent PSD, dispersion of coal and sorbent PSD, and the mean diameter ‘d50’ of the cyclone. To verify the model, real PSD of fly and bottom ash from the CFB boiler OFz450 were analyzed. A comparison is made in Fig. 16.13 of the measured and calculated PSD of fly and bottom ash. The boiler load was 100% at that moment. For both ashes, the calculated results agree well with the measurements. In Fig. 16.13, the PSDs of coal and sorbent are also presented. The solid lines represent the recommended coal and sorbent PSDs, and the points show real ones. For both coal and sorbent, one can also see divergences between real PSD curves and those required by the manufacturer. The solid mass flow, circulating in the boiler under such conditions (i.e., RFA = 72%, RBA = 42%) was calculated as about 250 kg/s. This means that the RR was about 25 at that time. The mass flow of recirculated fly ash was about 8 kg/s, and the mass flow of recirculated bed ash was about 1.2 kg/s. The cyclone mean diameter d50 was calculated as about 80 mm. The absence of ash recirculation and the decrease of boiler load changed the solid mass flow. This was calculated at about 100 kg/s (which means that RR equaled about 10). Lack of ash recirculation also caused the change in PSD dispersion of fly ash distribution because of worse separation conditions in the cyclone, due to the lower flue gas velocity. Under these conditions, the calculated d50 of the cyclone was about 90 mm.

© Woodhead Publishing Limited, 2013

Circulating fluidized bed combustion (CFBC)

733

Cummulative weight fraction (–)

1.0

RFA 72% 0.5

RBA 42%

Bottom ash-measurement Fly ash-measurement Bottom ash-calculation Fly ash-calculation Coal Required coal Required sorbent Sorbent

0.0 10

100

1000 d (mm)

10000

16.13 The comparison of measured and calculated PSD of fly and bottom ash.

The calculated value of RR = 25 is almost within the right area in Fig. 16.11. Nevertheless, it can be acceptable only for a medium LHV fuel. Such a RR value appears to be a critical one. The calculated influence of coal and sorbent PSD on RR value is shown in Fig. 16.14. The assumed change in the mean diameter d50 of sorbent was between 0.1 and 0.55 mm (the real d50 was 0.25 mm). The change for coal was between 0.4 and 1.4 mm (a real d50 was 1.1 mm). The decrease in d50 always causes an increase in RR and the total increase range is approximately the same for both materials. Based on the calculations, one can ascertain that if the sorbent’s mean diameter will decrease every 1 mm, the increase of total solid mass flow carried up in the chamber will increase about 0.08%. The same change of d50 for coal will result in a 50% change for the sorbent. The change of RR versus PSD dispersion is presented in Fig. 16.15. This change is of less importance. To analyze the curve for coal one can ascertain that the quality of milling, which is characterized by the parameter ‘s’, has no influence on the value of RR. This is clear, because the coal content in the bed is low due to its intensive combustion. The solution to increase the RR considerably is shown in Fig. 16.16. In this case, the relationship is strongly non-linear. One can determine that, as the cyclone’s d50 = 80 mm, the RR equals 25. The change of cyclone’s d50 down to d50 = 40 mm will cause a rapid increase of RR (up to almost 90).

© Woodhead Publishing Limited, 2013

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Fluidized bed technologies for near-zero emission combustion 30

RR (–)

20

Sorbent, s = 0.8 Coal, d50 = 1100 mm, s = 1 Coal, s = 1 Sorbent, d50 = 250 mm, s = 0.8 Cyclone d50 = 80 mm 10 0

200

400

600

800 d50 (mm)

1000

1200

1400

16.14 The RR versus mean diameter ‘d50’ of coal and sorbent. 30

RR (–)

20

Sorbent, d50 = 250 mm Coal, d50 = 1100 mm, s = 1 Coal, d50 = 1100 mm Sorbent, d50 = 250 mm, s = 0.8 Cyclone d50 = 80 mm 10 0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

s

16.15 The influence of coal and sorbent PSD dispersion on the RR.

From these calculations, such a value of RR should provide a wide possibility of temperature regulation in the boiler’s chamber. The new value of RR should also easily enable the unit to achieve its full output. Moreover, with such a RR value, neither RFA nor RBA will be necessary. The onsite improvement conducted for the cyclone in 1998 confirmed the above calculations. After the d50 of the cyclone was decreased, the boiler easily

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Circulating fluidized bed combustion (CFBC)

735

RFA = 0, RBA = 0 RFA = 30%, RBA = 0 RFA = 60%, RBA = 0 RFA = 0, RBA = 30%, 60%

RR (–)

100

Sorbent, d50 = 250 mm, s = 0.8 Coal, d50 = 1100 mm, s = 1 0

20

40

60 80 d50 of the cyclone (mm)

100

16.16 Influence of the cyclone’s mean diameter ‘d50‘ on the RR.

achieved the maximal output and the emission of SO2 and Nox became within the guaranteed limit. Nevertheless, the better cyclone efficiency will cause an increase in ash flow, which has to be drained from the bottom part of the boiler. As calculated, a change of cyclone’s d50 from 90 mm to 40 mm causes the increase in this ash stream from about 0.8 kg/s to 2 kg/s. The influences of RFA and RBA on the mass flow have also been presented in Fig. 16.16.

16.3.3 Selected aspects of heat transfer in the CFB boiler There are several CFBC models describing the heat transfer in the CFB boiler. in the present approach, experiences in modeling performed for a 670 t/h CFB boiler operated in Turow Power Station in Poland have been taken into account. The bed temperature profile is obtained from the energy balance for a steady-state condition according to the following equation (Krzywanski et al., 2010): (c p,g mg + c p,s ms ) ∂T ∂t = (c p,g mgU g + c p,s msU s ) ∂T + Qreact + Qair + Qrec – Qwalll – Qv – Qbott ∂z [16.37]

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Fluidized bed technologies for near-zero emission combustion

736

where cp,g is specific heat capacity of gas, J/kg◊K; mg is mass of gas, kg; cp,s is specific heat capacity of solids, J/kg◊K; ms is mass of solids, kg; T is temperature, K; t is time, s; Ug is flue gas velocity, m/s; Us is particle velocity, m/s; z is distance of cross-section area from the grid, m; Qreact is energy of reactions, W; Qair is energy of air, W; Qrec is energy of recycled particles, W; Qwall is energy lost through walls, W; Qv is energy used for moisture vaporization, W; and Qbott is energy of bottom ash, W. in the determination of heat transfer to the surfaces of the following heat exchangers incorporated in the combustion chamber, relations were used, which indicate the heat transfer coefficients for the membrane walls and superheater surfaces, while taking into account conduction, convection and radiation (Diego et al., 1996). The thermal computations carried out for 670 t/h CFB boiler were based on relations which describe the heat-transfer conditions of the dense zone and the dilute zone of the combustion chamber, separated by the level of the secondary air feed points. The overall heattransfer coefficient can be calculated from the following relations (Diego et al., 1996): ∑

For the membrane wall surfaces: – in the dense zone: h = a1 –

in the dilute zone: h=



[16.38]

a1 (1 + e f a1)

[16.39]

for the SH I and SH II: h=

a1 Ê e f + 1ˆ 1+Á a1 Ë a 2 ˜¯

[16.40]

where ef is the fouling coefficient; ef = 0.001–0.003 m2k/W. The coefficient of heat transfer from the bed to the heat exchangers surfaces considers two mechanisms, i.e. convection and radiation: ∑

in the dense zone: a1 = (aconv + arad)



[16.41]

in the dilute zone: a1 = x (aconv + arad)

[16.42]

where x = 0.85 is the coefficient of exploitation, which takes into account the non-uniform flow of flue gas in a short distance to the heat exchangers surfaces.

© Woodhead Publishing Limited, 2013

Circulating fluidized bed combustion (CFBC)

737

The convection heat transfer coefficient considers the convection of flue gas and the convection of bed material particles: aconv = ag + ap

[16.43]

In the computation of the coefficient of convection heat transfer to the heat exchangers surfaces, the following relations can be used: ∑

in the dense zone: aconv = a(rp)b



[16.44]

where a = 13.7; b = 0.33; and rp is density of bed material, kg/m3. in the dilute zone:

a conv = 0.0214

0.8 ˘ È Ê d ˆ 0.67 ˘ lc ÈÊ wde ˆ – 100˙ Pr 00..4 Í1 + Á e ˜ ˙ Í Á ˜ Ël¯ ˙ de ÍË v ¯ Î ˚˙ ÎÍ ˚

[16.45]

where lc is thermal conductivity for flue gas, W◊m–1◊k–1; w is average flue gas velocity, m◊s–1; v is coefficient of kinematic viscosity for flue gas, m2◊s–1; and l is height of the membrane wall, m. The equivalent diameter of the combustion chamber can be calculated based on the following expression de = 4A O furn

[16.46]

where A is boiler cross-section area, m2; and Ofurn is combustion chamber perimeter, m. The coefficient of heat transfer by radiation can be determined from the following expressions: ∑

in the dense zone:

a rad = s 0 (T p2 + Tw2 )(T p + Tw )

1 (e w–1 + e –1 p – 1)

[16.47]

where Tp is fluidized bed temperature, K; Tw is membrane wall temperature, K; s0 is the Stefan–Boltzmann constant 5.67◊10–8, W◊m–2◊k–4; and ew = 0.8 for the membrane wall, and: 4

Ê Tw ˆ e p = e a0.4 (1 – Y ) Á ˜ + Y Ë Tp ¯

[16.48]

where ea = 0.7–0.81 for the bottom ash, Y = 1 – e–0.16Ar

0.26

© Woodhead Publishing Limited, 2013

[16.49]

738



Fluidized bed technologies for near-zero emission combustion

in the dilute zone:

a rad

ÊT ˆ 1 – Á wz ˜ Ë Tg ¯ a +1 = 5.698 ·10 –8 w asTg3 2 T 1 – wz Tg

4

[16.50]

where aw is emissivity of heat exchanger surface, –; as is emissivity of gas-solid phase, –; Twz is temperature of the dirty wall, K; and Tg is average flue gas temperature, K. emissivity of the flue gas can be calculated based on the following expression: as = 1 – et

[16.51]

where t is optical thickness which is the sum of the optical thickness of the flue gas tg and optical thickness of gas–solid phase tp. t = tg + tp

[16.52]

The optical thickness of the lean phase is described by the following relation: Ê ˆ Á ˜ b2 4.1 t p = 2 2 Á1 – 3˜ (d pT ) Á 1 + 30 · 10 ˜ (r* · s ) ¯ Ë

[16.53]

where s is thickness of radiating layer, m; b2 = 0.6–0.7, and:

r* =

kby Vg

[16.54]

where kb is bed material circulation ratio, –; y is coefficient considering the internal circulation of bed material in the combustion chamber, –; and Vg is the real volume of a flue gas for the combustion air factor, m3/kg. The thickness of the radiating layer is described with the functions: ∑

for the dense zone: s = 3.6



V furn Aw

[16.55]

where Vfurn is combustion chamber volume, m3; and Aw is overall surface area of walls in combustion chamber, m2. for the dilute zone, where SH I and SH II are located:

© Woodhead Publishing Limited, 2013

Circulating fluidized bed combustion (CFBC)

s=

A free

3.6 V furn + ASH SHn n + AS SH H

V free ˆ Ê ASH · ÁË1 + A free + ASH ˜ SH SHn Hn n V furn ¯

739

[16.56]

where Afree is surface area of free space inside the combustion chamber, m2; ASHn is surface area of walls combined to the superheater, m2; ASH is superheater surface area, m2; and Vfree is volume of free space inside the combustion chamber, m3.

16.3.4 Separators in CFB boilers Many published papers have described procedures or provided information relevant to the estimation of the pressure drop and separation efficiency of a cyclone. However, one should keep in mind that the cyclone performance in a CFB boiler is not the same as in classical cyclones, as a result of the comparatively higher inlet solids loading found in cyclones when used in a CFB boiler. Gas–solids separators used in CFB boilers can be divided into the following main categories: ∑ ∑ ∑ ∑ ∑

hot cyclones in which internal area is covered with refractories, compact separators made of water/steam-cooled membranes (Foster Wheeler type), separator inside the combustion chamber (CYMiC type by kvaerner), cyclones made out of water/steam-cooled membranes, impact separators without the use of centrifugal force (Babcock & Wilcox U-beam separator).

The schematic diagram of the CFB separator is shown in Fig. 16.17. The cyclone is located at the outlet of the combustion chamber and it separates particles from flue gases recycling them to the lower part of the combustion chamber. The cyclone is designed for a very high separation efficiency. In the steam-cooled separator or cyclone, the tubes forming the cyclone are integrated with the water–steam circuit. The cyclone is an extension of the furnace pressure parts. as a result, the entire separator enclosure is steam cooled. The interior area is covered by a thin refractory, resulting in a significant reduction of weight. Because the internal walls are covered with thin refractories, the start-up time is significantly reduced as compared to hot cyclones. Due to the cooled construction of the steam-cooled cyclone, there is no need for heavy refractory lining. This minimizes maintenance and allows for shorter start-up times since the heat capacity of refractory materials limits the heating rate during the start-up (otherwise the refractory can be damaged). High efficiency of the cyclone provides proper solids recirculation, which in turn facilitates uniform temperature distribution in the CFB furnace. There

© Woodhead Publishing Limited, 2013

740

Fluidized bed technologies for near-zero emission combustion

16.17 The new CFB cyclone separator designed by Rafako S.A.

is a substantial increase in the solids recirculation rate with an increase in separator efficiency. The most important parameters are the separation efficiency and the pressure drop. The overall separation efficiency is simply calculated as the mass fraction of the feed solids captured by the cyclone. Overall efficiency is usually used in CFB boilers. However, this is not a good measure for characterizing the intrinsic separation performance of the cyclone, since it depends not only on the cyclone but also on the size and solids density. Therefore, so-called ‘grade efficiency’ is used, which is the separation efficiency for a given feed particle size or range of particle sizes. ‘Grade efficiency’ can be defined as a measure of the percentage of particles of a specific size that enter the separator and are separated. For example, if a million particles of 100 mm enter the separator and 900,000 are retained, then the grade efficiency of the separator would be 90% for 100 mm particles. The total separation efficiency in CFB boilers is the sum of the two separation efficiencies (Muschelknautz and Trefz, 1993; Muschelknautz et al., 1994):

© Woodhead Publishing Limited, 2013

Circulating fluidized bed combustion (CFBC)

htot = hlim + hi = 1 –

mlim mlim + R (d *) me me Ai i

741

[16.57]

where separation efficiency due to inner vortex is given as (Muschelknautz and Trefz, 1993; Muschelknautz et al., 1994): È Ê d * · 0.71/nAi ˆ nAi ˘ mlim * hi ª R (d ) = exp Í– Á i ˜¯ ˙ me Ai i de* ÍÎ Ë ˙˚

[16.58]

separation efficiency due to the wall is (Muschelknautz and Trefz, 1993; Muschelknautz et al., 1994):

hlim = 1 –

mlim me

[16.59]

and limited loading ratio is:

mlim = 0.025

de* (10 me )k d50

[16.60]

The size of the cyclone (or of a number of cyclones) is defined by the gas volume flow. Larger capacity CFB boilers will have a higher gas volume flow and so will require a larger cyclone, or a greater number of cyclones when using a ‘modular’ approach. The proper level of the solids flux depends on the efficiency of the separator, expressed as cut size of d50. This dependence has a strong nonlinear character. Even a small increase in separator efficiency improves heat exchange conditions and the reduction of temperature in the lower part of the furnace. intensive research on cyclones resulted in a reduction of separator cut size (d50) from 150 mm to 80 mm, resulting in an increase in solids circulation rate (RR – mass stream of circulating solids per mass of fuel fed into the chamber). This allows reserves of boiler control in the furnace and an improvement of heat exchange conditions. Decreasing d50 of the separator increases the solids recirculation to a level which allows stable boiler operation, even with very low ash content. Similar tendencies occur in Compact separators, which have slightly lower solids separation efficiencies than a cyclone. Although conventional cyclone diameters are large (up to 10 m, 235 MWe Turów CFB), in practice they achieve high efficiency as a result of the influence between solid particles in the solids flux at the inlet to the cyclone. Good circulating material consists of particles that should be retained in the circulating bed between 50 and 350 mm. The mean particle size of the circulating material should be within the range of between about 180 mm and about 250 mm. Poor collection efficiency of the separator results in: ∑

loss of fine particles from the circulating loop,

© Woodhead Publishing Limited, 2013

742

∑ ∑

Fluidized bed technologies for near-zero emission combustion

coarse circulating material which is more difficult to entrain higher in the furnace, and consequent poor thermal transfer by particle convection.

One means of increasing separator efficiency is the proper selection of the inlet channel to the separator to achieve a high mass density, which is directed to the wall of the separator. The effect of such a modification is shown in Fig. 16.18. There is an improvement of fine particles separation (under 20 mm), giving the opportunity to utilize fine limestone and make-up sand. The curve shown in Fig. 16.18 is similar to that for a granular bed filter. The separation mechanism for high solids loads in the cyclone is presented in attached films from studies of the separation efficiency at 235 MWe CFB boilers. The effect of higher separation efficiency on solids recirculation rate (RR) with different ash content fuels is shown in Fig. 16.19. The circulating material is initially made up of sand. Once solid fuel combustion commences, ash and limestone also form part of the circulating material. Coal ash helpfully contributes to the circulating bed but biomass ash does not. If the separation efficiency of the separator/cyclone is adequate, a satisfactory bed can be maintained when firing coal without the need for significant portions of make-up sand. If the separator is under-performing, too many of the optimally sized particles required in the circulating loop will escape into the back pass. This will result in:

1.0

Separation efficiency hi (–)

0.8

0.6

0.4

0.2

After modification Before modification

0.0



1

10 100 Particle size di (mm)

16.18 Separation efficiency as a function of particle size.

© Woodhead Publishing Limited, 2013

1000

Circulating fluidized bed combustion (CFBC)

743

250 Ash content in coal 17% Ash content in coal 11%

200

Ash content in coal 6%

RR

150

100

50

0

20

40

60 80 100 120 d50 of the cyclone (mm)

140

160

16.19 Influence of d50 of the separator on the solids recirculation rate (RR) depending on ash content in fuel.



difficulties achieving the required levels of heat transfer in the furnace, ∑ high bed temperatures, ∑ higher consumption of make-up sand, ∑ erosion in the back pass, and ∑ reduced environmental friendliness. The pressure drop over the CFB cyclone is subdivided into the following contributions: ∑

Losses in the entry i.e. contracting pressure drop produced from the contraction from the large area of the combustion chamber to a smaller area of the cyclone inlet; this pressure drop depends on the ratio of inlet cyclone area to the combustion chamber area and indeed on the gas inlet velocity. The acceleration of solids pressure drop (one of the largest individual pressure drops) depends on the particle velocity and the solids loading. ∑ Losses in the separation space (the so-called barrel friction pressure drop and the gas reverse pressure drop). The barrel pressure drop is a function of the hydraulic diameter of the cyclone (hydraulic diameter = 4 (inlet area)/inlet perimeter). ∑ Losses in the vortex finder (this pressure drop is the largest); this pressure drop depends on the relative dimensions, specifically the square of the ratio of the cyclone diameter to vortex tube diameter. This pressure drop

© Woodhead Publishing Limited, 2013

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Fluidized bed technologies for near-zero emission combustion

can be reduced by using a conical vortex finder (Hoffman and Stein, 2002). The relative dimensions of the cyclone have a significant effect on the pressure drop. Indeed, smaller inlet height gives higher pressure drop in the entry. However, the mechanism of solids flow from the inlet through the vortex finder is more complicated than that from classical cyclones. Additionally, the relative dimensions of the cyclone strongly influence its separation efficiency (although the separation efficiency also depends, in part, on the particle size, solids loading and PSD). Ratios such as the height of the cyclone body to cyclone diameter and dust exit diameter to cyclone diameter are the most important relative dimensions that affect the separation efficiency. There is evidence amongst published data that the relative dimensions such as inlet duct dimensions (ratio of inlet height to inlet width and inlet configuration), cyclone length to cyclone diameter, vortex finder area to inlet area and many others play a significant role in improving separation efficiency. Figure 16.20 and Table 16.4 illustrate some information on typical proportions of cyclones, including CFB cyclones. However, the proper design of CFB separators requires experience and knowledge, which can only be obtained by operating CFB plants and be refined through an extensive process of research, testing and development. The normal practice in CFB design is to build one or more larger diameter cyclones or separators. Smaller cyclones have higher separation efficiency but it is difficult to put refractory in such small cyclones. As the size of CFB boiler increases, the mass flow of gas and solids entering the separator increases proportionally. To accommodate high gas–solids flow, one can m

L F

K Dc

S H B

E

16.20 Standard proportions of cyclones.

© Woodhead Publishing Limited, 2013

Circulating fluidized bed combustion (CFBC)

745

Table 16.4 Typical proportions of vertical reverse-flow cyclones (see Fig. 16.20)

General CFB

K DC

L DC

m DC

F DC

S DC

H DC

E DC

References

0.80

0.350

0.45

0.5–1.0 1.00 2.00 –

High throughput 0.75

0.375

0.75

0.875

1.50 4.00 0.375 Stairmand (1951)

0.80

0.350

0.75

0.850

1.70 3.70 0.400 Swift (1969)

General purpose 0.50

0.250

0.50

0.600

1.75 3.75 0.400 Swift (1969)

0.50

0.250

0.50

0.625

2.00 4.00 0.250 Lapple (1951)

0.44

0.210

0.40

0.500

1.40 3.90 0.400 Swift (1969)

0.50

0.220

0.50

0.500

1.50 4.00 0.375 Stairmand (1951)

High efficiency

16.21 Foster Wheeler’s Compact separator design.

design the boiler with several small cyclones instead of a few larger ones. The performances of the separator, furnace, and returned leg impact the temperature profile in the combustion chamber and other critical process characteristics. However, all of these components work together to meet steam generation rate, boiler efficiency and emission levels. Foster Wheeler’s original Compact separator design (as shown in Fig. 16.21) was rectangular in cross section and had flat, straight panel walls, rather than the curved walls typical of the conventional cyclone separator.

© Woodhead Publishing Limited, 2013

746

Fluidized bed technologies for near-zero emission combustion Coarse particles

Fines

Impact

Gas

Separation

16.22 Mechanism of impact separation.

Flow pattern in the Compact separator is almost the same as in a cyclone. There is no fundamental difference between the separator and large cyclones with respect to the separation efficiency. The gas flow in the separator is a swirling vortex type, despite the rectangular cross section of the separator. Due to the centrifugal effect, solid particles are swept to the separator walls and continue to flow along the wall surface. In the corners of the separator, the particles form dense concentrations. The Compact separator does not operate on the impact principle, as illustrated in Fig. 16.22, which is used in U-beam B&W CFB boilers. Inlet flow is not impacting the rear wall (which would have resulted in a curtain of solids along the rear wall that would trap other fine solids impacting the curtain of falling ash). The hexagonal Compact separator used in the 460 MWe CFB boiler in Lagisza, Poland constitutes a significant improvement on the Compact separator as compared to the standard rectangular design. The result of these improvements is that more and finer particles will circulate in the CFB loop (furnace, separator, return leg-comer and loop seal), leading to improved process performance.

16.4

Reliability and availability of CFB boilers

Boiler availability is a major concern of any boilers. Generally, the guaranteed availability of the CFB boiler should be no less than 95%. The most common measures for CFB boiler performance are availability factor (AF), forced outage factor (FOF) and serviced outage factor (SOF) defined as follows: AF =

∑in=1 t pi + t ri · 100% ∑in=1 t ki

FOF =

∑in=1 t ai · 100% ∑in=1 t ki

© Woodhead Publishing Limited, 2013

[16.61]

[16.62]

Circulating fluidized bed combustion (CFBC)

SOF =

∑in=1 (t k + t s + tb )i · 100% ∑in=1 t ki

747

[16.63]

Disturbances that are caused by external reasons, not under the control of plant management, are not classified as unavailability. As an example, the availability of CFB units No. 1–3 (670 MWth with hot cyclones) operated in Turow Power Plant, Poland will be presented. The design availability of CFB units No. 1–3 was settled at the level of 85%. Problems in achieving this value appeared from the beginning of operation of these units (Fig. 16.23(a)). This value was achieved in 2003 and 2004. The main influence on the availability of CFB units was the maintenance time. Faults resulting from project and assembly defects were resolved in the first years. As can be seen in Fig. 16.23(a), it was possible to decrease the SOF in the fourth year of operation, below 10%. However, this factor began to increase again from 2004 as a result of unplanned current repair, despite the fact that average repairs were carried out in these years. The main reason for low availability of CFB units No. 1–3 was boiler breakdowns (Fig. 16.23(b)). The erosion of refractories and heat surfaces are the most essential problems of CFB boiler performance. in addition, disturbances began to appear in turbine performance in 2003 resulting in disconnection of CFB units. Defects in the transformers also influenced the high unavailability of these units during their usage. in spite of enlarging the expenditure on scheduled maintenance, including current repairs, the forced outage factor from the beginning of exploitation remained at a high level of about 5%. The first years of operation resulted in among others: improvement in the durability of compensators and air nozzles, modernization of the convective path, erosion protection for the bottom part of membrane walls, modernization of the tubular air heaters, exchange of refractories to a more durable one. Later problems were also revealed with the dynamic state of turbines. Similar to CFB units 1–3, the initial period of operation of CFB Compact units No. 4–6 was characterized by high FOF and lower than planned availability factor AF (Fig. 16.24(a)). A high decrease of FOF below 2% already in the second year of operation and its stabilization at values of 1% and AF remaining at a level above 90% testifies to the inevitable transition of the CFB Compact units from the running-in period until their normal operation. Enlarged time of planned repairs in the first three years of boiler operation and the high FOF in the first year was the result of the need to remedy the faults of the boilers and turbines and their auxiliary devices (Fig. 16.24(b)). Low shutdown times testify that these problems were insignificant. in order to improve plant availability, there is a need for preventive maintenance, regular monitoring and inspection of the plant. The CFB boiler

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100 FOF SOF 90

AF

15

80

10

70

5

60

AF

FOF, SOF

20

0

50 1999

2000 2001

2002 2003 2004 2005 Year of operation (a)

2006

700

2007

Other Turbine Generator Boiler

600

Shutdown time (h)

500 400 300 200 100 0 1999

2000

2001

2002 2003 2004 Year of operation (b)

2005

2006

2007

16.23 Performance factors (a) and forced outages (b) of CFB units No. 1–3 with hot cyclones.

is a complex system that requires proper monitoring and maintenance. Based on data gathered from CFB boilers, it can be said that CFB boilers have similar availability and maintenance to pulverized boilers.

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749 100

FOF SOF

90

AF

15

80

10

70

5

60

AF

FOF, SOF

20

50

0

2003

2004

2005 2006 Year of operation (a)

2007

800 Other Turbine Generator

Shutdown time (h)

600

Boiler

400

200

0

2003

2004

2005 2006 Year of operation (b)

2007

16.24 Performance factors (a) and forced outages (b) of CFB Compact units No. 4–6.

16.4.1 Selected operational problems of large-scale CFB boilers During operation of some CFB boilers, there are obstacles to the boiler operating properly. Common disturbances in CFB boilers burning coal are:

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fan failures, fuel feed blockages, bed sintering, damage to refractories and heat surfaces, erosion in furnaces and backpass, expansion joint damage and also to air and flue gas duct.

If the input materials (sand, fuel, limestone) are out-of-specification and, as a result, the particle size distribution is also far from the design values, then serious problems can be expected with the operation of the CFB boiler, including: higher temperatures, higher emissions, sintering, lower capacity of the boiler, large differences between the upper and lower part of combustion chamber, etc. The out-of-specification material negatively affects hydrodynamic properties in the combustion chamber, temperature distribution and heat transfer, and promotes excessive erosion. The quality of make-up sand is particularly important when burning biomass with low ash content, coal alone without limestone, or coal when burning for a relatively short period. When make-up sand particles are too coarse, higher bed temperatures result due to lower heat transfer in the combustion chamber. Coarse make-up bed material results in coarser circulating material, poor solids flow and suspension density in the upper part of the combustion chamber, which in turn has a negative impact on: heat transfer, bed temperature, sintering, combustion process and emissions. It is also very important to control the bed quality through the bottom ash removal system. The draining of the bottom ash depends on the bed pressure. Usually, the bed pressure is adjusted within certain limits when the draining automatically starts and stops. The bottom ash analysis should be done frequently to maintain the bed quality. During the combustion of alternative fuels, the majority of agglomerates are created due to the existence of a liquid phase consisting of alkalis (Na, Ca, K, etc.), CaSO4 or aluminasilicates. If the fuel consists of a large amount of such components, the agglomeration may bring about difficulties in proper operation of the furnace. Based on research of agglomeration in fluidized bed boilers, it was found that the knowledge of properties of any particular fuel and its ash, as well as sorbent used for desulfurization, is crucial to avoid agglomeration; particularly that the mechanism of agglomerate formation depends on fuel type. Defluidization may sometimes occur at temperatures lower than ash softening temperature, due to the existence of local liquid spots. Accordingly, attention should be paid to any minor element in ash and sorbent. Apart from chemical composition, agglomerate formation is also affected by combustion conditions, particularly oxygen concentration. For low concentrations of O2, the metal oxides are reduced and the pure metal produced has a lower melting temperature. Avoiding agglomeration is in general possible by keeping a uniform temperature in the furnace, and

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avoiding changes of fuel or sorbent (e.g. from limestone to dolomite) or change of particle size distribution of solids (e.g. from coarser to finer ones). Agglomerates in fluidized bed combustors are usually created in those areas where temperature is higher than the ash softening temperature. This may be due to poor fluidization and mixing or low gas velocity in the bed. Gas velocity at which defluidization occurs increases linearly with bed temperature. However, fluidization without agglomeration at temperatures higher than the initial temperature of agglomerate formation is possible, but higher gas velocity is required. Experience obtained from CFBC of various fuels led to the conclusion that the most important factors in preventing agglomeration are the chemical composition of fuel, ash and the bed material. in the case of burning biomass in the CFB boiler, higher alkali content causes fouling and sintering of the bed and it is sometimes necessary to ‘clean’ or flush the bed by increasing sand consumption. When burning biomass, make-up material of the circulating bed is formed mainly by: ∑ ∑ ∑

classified sand (size distribution is proper), ash from biomass, unclassified material (foreign sand, gravel, metal).

The amount of foreign sand in the biomass and its quality are extremely important since ash content in the biomass is low and the fluidization process is calculated based on classified sand particles. Unclassified material can also cause serious problems such as erosion in the bottom part of the combustion chamber. Fine sand can participate in circulation through the CFB loop and can create erosion in the upper parts of the boiler. Fouling depends on ash-forming elements in fuel which react with the flue gas and form fouling components. In order to estimate fuel tendency to fouling, it is useful to calculate the fouling index (Fu): Fu =

Fe2O3 + C CaO aO + MgO O+N Naa2O + K 2O + P2O5 · ((Na Na2O + K 2O ) SiO2 + Al2O3 + T TiO iO2 [16.64]

When the Fu ≤ 0.6, no fouling occurs in the boiler. When 0.6 ≤ Fu ≤ 40, moderate fouling can be observed, but when Fu ≥ 40, very high fouling occurs in the CFB boiler. oxide analysis of two Polish coal (hard coal and lignite) ashes and two biomass (two specimens of straw) ashes is given in Table 16.5 (Rajczyk and Bednarek, 2011). As can be seen from Table 16.5, both hard coal and lignite have low tendency to fouling. Straw is usually considered as a problematic fuel. However the Fu factor of straw a is also low (1.46) and similar to that of coals. On the other hand the Fu index of straw B is far higher (165.5) making straw B the fuel with an extreme tendency to fouling. it should be concluded that fouling tendency is a parameter of an

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Table 16.5 Oxide analysis of ash and Fu index of main fuels burned in Polish CFB boilers

Fuel

Oxide Fe2O3 CaO

Hard coal Lignite Straw A Straw B

5.1 5.00 1.35 0.35

8.7 2.22 8.74 11.8

MgO

Na2O K2O

P 2O 5

SiO2

Al2O3 TiO2

5.1 2.25 2.48 3.79

1.5 1.18 0.25 0.96

0.2 0.35 2.38 3.81

39.2 54.2 73.1 16.7

25.3 28.7 2.36 0.38

2.1 2.86 5.16 43.3

Fu

1.5 1.25 1.17 0.65 0.17 1.46 0.03 165.5

individual fuel and depends on chemical composition, mostly on Na and K content. Fouling tendency of fuel may also be predicted using characteristic temperatures from ash analysis and sintering analysis. Erosion is one of the major concerns for any CFB boiler operators. However, erosion caused by flow of solids particles happens inside every coal-fired boiler. Materials are required to withstand high temperatures and high velocity of particles. This concern had been alleviated in the CFB boiler by adopting the philosophy of eliminating any unnecessary discontinuities that would change the direction of downward and upward flowing bed particles. In locations where such discontinuities must exist, erosion protection is designed into the systems with shielding, or weld overlay, or refractory coverage, or a combination of protection methods. The particles in the lower combustion chamber are very abrasive. The refractory starts from the grid floor to the transition from the combustion chamber vertical wall to the sloped wall. Experience of some boiler manufacturers has shown that erosion of watercooled walls in this area can be avoided by using, for example, kick-out pipes. Any welding on the membrane wall of the combustion chamber must be performed in such a way that will not result in vortex formation, which will lead to excessive erosion of the tubes. The surface of the weld should be very smooth with no protrusions from the membrane wall surface. If some pieces of the welding stand proud of the tubes, then particles flowing down strike the weld surface, resulting in high erosion. Thermal or plasma spray coatings are one of the best solutions to protect tubes against erosion. Some foreign objects accumulate at the bottom of the bed causing local erosion. This is not unusual on a CFB boiler. Foreign objects such as metal, steel plates, etc., found in the CFB boiler can seriously damage the feeding system and shut down the boiler. Furthermore, unclassified material can be a cause of grid blocking. For smooth operation of the CFB boiler, the grid nozzles should be carefully cleaned before the start-up of the boiler. Proper bed maintenance and bed cleaning are necessary. The return leg and the grid nozzles are the most important parts in the CFB boilers. If some of the nozzles are blocked, then the division between primary air and secondary air changes.

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Development strategy and challenges of CFBC technology

Currently, major factors determining the development of CFB technology include: standards of gaseous emissions (LCP Directive), unlimited access to various types of fuel and strong competition in the field of new energy technologies (clean coal technologies). The last 20 years have seen rapid development of the technology of fuel combustion in a circulating fluidized bed, which consisted of progressively increasing the size of the units. This trend is noticeable in the case of the development of fluidized bed boilers of the first and the second type of CFB generation. Depending on the region of the world, CFB technology has become more or less widespread. Since the 1990s, there can be observed a fast increase in installed power units based on CFB technology, particularly in North America and Asia. In Europe, 63% of electricity production takes place in more than 20-year-old CFB boilers burning coal. At the moment, the world’s largest fluidized bed combustion boiler is the second generation of once-through supercritical circulating fluidized bed (CFB OTU-SC) with a capacity of 460 MWe, which is installed in the TAURON Wytwarzanie S.A., a branch of Lagisza Power Plant, Poland. The schematic diagram of the 460 MWe OTU-SC CFB boiler is shown in Fig. 16.25, while basic construction parameters are shown in the Table 16.6. The CFB boiler operated in Lagisza Power Plant is designed to burn bituminous coal as a main fuel. The fuel properties for Lagisza’s boiler are shown in Table 16.7. As can be seen from Table 16.7, there is a possibility for combustion of additional fuel in the Lagisza CFB boiler, in particular coal slurry. Coal slurry can be combusted with maximum 30% share by fuel heat input. Light oil is used as a start-up fuel. Unlike typical CFB boilers, this boiler is equipped with Benson vertical tubes which is a new supercritical steam technology. On the grid level, the combustion chamber has a cross-section diameter 27.6 ¥ 5.3 m (grid width ¥ grid depth). On the height 8.95 m from the grid, the combustion chamber cross-section area has a diameter of 27.6 ¥ 10.6 m (width ¥ depth). Total height of the combustion chamber is 48 m. On the bottom of the combustion chamber is the fluidization grid with arrowhead nozzles of primary air. Secondary air nozzles are positioned at 2 m, 4 m and 6 m from the grid (SA right wall and SA left wall). The boiler is equipped with eight compact separators of a loose material, four located along each side wall. The cross section of the separators is octagonal. The separators are made of membrane walls, which are cased with a thin layer of erosion-resistant refractory. A similar construction is used for the solids particles’ return leg to the INTREX chamber, which is integrated with boiler’s side walls. The INTREX chamber has a wall seal structure. In the INTREX chambers, the heat is transferred with circulating material to the INTREX superheater and

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Separator outlet duct

Crossover duct

Separator outlet duct

Crossover duct

Furnace

Compact separator (Typ.-4)

Compact separator (Typ.-4) Full height evaporator wingwalls (Typ.-3) 38¢–0¢¢ depth 129¢-8¢¢ width

Fuel feed chain conveyors Fuel feed chute (Typ.-8)

Fuel feed chain conveyors

Wall Seal (Typ.-8) Intrex Liftleg (Typ.-8)

Fuel feed chute (Typ.-8)

Intrex Bypass (Typ.-8)

Intrex SH-IVb

Intrex SH-IVa Intrex SH-III

Secondary air fan (Typ.-2)

Ash conveyor

Intrex RH-III

Ash screw coolers S/C screw conveyor

Primary air fan (Typ.-2)

16.25 Schematic diagram of the 460 MWe OTU-SC CFB boiler – front view (Goidich et al., 2000).

Table 16.6 Design parameters of 460 MWe OTU-SC CFB boiler in TAURON Wytwarzanie S.A., a branch of Lagisza Power Plant Parameter

Unit

Data

Capacity Net electrical efficiency Boiler type Main steam flow Main steam pressure Main steam temperature Reheat steam flow Reheat steam pressure Reheat steam temperature Feed water temperature

MWe – – kg/s MPa °C kg/s MPa °C °C

460 44 OTU-SC 360 27.5 560 307 5.5 580 289.7

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Table 16.7 Fuel properties for the Lagisza CFB boiler

Lower heat value Moisture content Total ash content Sulfur content

MJ/kg % % %

Coal range

Coal slurry range (max 30% input)

18–23 6–23 10–25 0.6–1.4

7–17 27–45 28–65 0.6–1.6

INTREX reheater. In the loop seal of the combustion chamber, there is a lack of heat exchangers for heat removal, but there are located points of solid particles’ returns from INTREX chamber (siphoned type of construction with rectangular cross section), points of fuel and sorbent (limestone) feeding, make-up sand and points of recirculation fly ash feeding. The CFB boiler at the TAURON Wytwarzanie S.A. Lagisza Power Plant shows off the main advantages of a second generation CFB techology and produces standard emissions that meet permissible concentrations of SO2 in flue gases. For all units, the concentration of SO2 was lower than 200 mg/m3n. Concentration of this component was contained in the range from 73 mg/mn3 to 200 mg/mn3. Nitric oxides concentrations from 131 mg/mn3 to 198 mg/mn3 were observed. In the case of CO levels, the minimum concentration was 24 mg/mn3 and the maximum 100 mg/mn3. Dust emission was from 5 mg/mn3 to 27 mg/mn3 and complied with appropriate environmental regulations for solid particles. Current emission levels for all gases, pollution and dust meet permissible values at the TAURON Wytwarzanie S.A. Lagisza Power Plant. The Korea Southern Power Co., Ltd (KOSP) is building the largest and most advanced supercritical circulating fluidized bed steam generators ever, to power their Samcheok Green Power Plant. The facility is located in Samcheok, within the Gangwon province of Korea. Not only is Samcheok the first supercritical CFB plant ordered in Korea, but it is also the third CFB ordered in the world that utilizes advanced vertical-tube, once-through supercritical steam technology. Table 16.8 presents the parameters of the CFB boilers. For modern power engineering technologies to be able to meet ever stringent environmental requirements, especially in the aspect of CO2, new concepts of the power plant are being developed, in which innovative solutions are used, such as oxy-fuel combustion in CFB boilers. The problems of combustion in oxygen-enriched atmospheres have been given increasingly large coverage in the relevant literature recently (Czakiert et al., 2006, 2010; Tan et al., 2012; Eddings et al., 2009; Zheng, 2011). The high interest enjoyed by the oxy-combustion process is primarily due to its key advantages, such as the increased energy conversion efficiency and the possibility of direct CO2 sequestration. Integral to the implementation of oxygen-enriched combustion

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Fluidized bed technologies for near-zero emission combustion Table 16.8 Parameters of CFB boilers at Samcheok Power Station, Korea Parameter

Unit

Data

Plant electrical output (gross/net) Unit steam flow (SH/RH) Steam pressure (SH/RH) Steam temperature (SH/RH) Feedwater temperature Furnace width Furnace depth Furnace hight Fuel

MWe kg/s bar °C °C m m m –

4 ¥ 550/4 ¥ 433 MWe 437/356 257/53 603/603 297 39.5 10.7 52 International coal, biomass

are the gas separation processes. This concerns especially the carrying out of the preliminary separation of atmospheric oxygen and nitrogen from the air supplied to the combustion chamber, which results in either a partial or complete elimination of N2 from the combustion process. With certain solutions, the separation of combustion gases of a considerably increased CO2 fraction (compared to the CO2 concentration at a level of 15%, being typical of the conventional combustion process) also becomes necessary. The concept of oxygen-enriched combustion means that there is no need for fuel recovery; the boiler is supplied with a gas mixture, whose oxygen concentration is higher than the air O2 concentration. The flue gas from an oxy-combustion is essentially CO2 and H2O along with nitrogen, oxygen, and trace gases like SO2 and NOx. Thus, CO2 can be relatively easily captured and stored. Another unique advantage of oxy-combustion is the complete avoidance of NOx emissions, usually generated by high temperature reaction in conventional air combustion. Indeed, NOx emissions are much lower (typically 25–50% lower than for air-firing) during O2/CO2 combustion since part of the recycled NO is reduced to molecular nitrogen. Summarizing, the main benefits of the O2/CO2 coal combustion are as follows: ∑ ∑ ∑ ∑ ∑ ∑ ∑

efficient recovery of CO2 from flue gases, high thermal efficiency due to reduced volume of flue gases by approximately 75%, ability to minimize unburned carbon, flue gas treatment system can be made compact and reduced by approximately 75%, lower energy loss for gas cleaning and separation, NOx and SOx emissions can be reduced, the higher temperatures increase the radiation (the flame emissivity is higher due to the higher concentration of CO2 and H2O) and consequently the heat transfer and processing rates.

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On the other hand, the following problems remain to be solved: ∑

∑ ∑ ∑ ∑ ∑ ∑

much power is required for oxygen production resulting in a decrease in the total power efficiency resulting in high operating costs; however, the cost of oxygen is expected to decrease due to improvements in oxygen production, recovered CO2 contains impurities, due to the recycling of the flue gases into the boiler, corrosive components play a significant role, the intensified heat output can cause refractory damage and overheating, higher oxygen concentration in the feed gas into the boiler leads to higher temperature causing potential slagging and agglomeration problems, intensifying the process with OEC can adversely affect the heat transfer distribution in the combustion chamber, no commercial oxygen combustion power plants are operating today, due mainly to the high cost of producing oxygen; however, pilot-scale studies have demonstrated that there are no significant technical barriers to oxy-combustion in CFB boilers.

The main challenge in the further development of CFB technology is building the demonstration oxy-fuel plant within the Compostilla Project named ‘OxyCFB300’, in Spain. The capacity of the CFB boiler is to be 323 MWe.

16.6

Conclusion

The use of a circulating fluidized bed for power generation is a rapid growing technology in the world. The ability of CFBs to burn a wide variety of fuels while meeting strict emission control regulations makes them an ideal choice for burning such fuels as high sulfur coal, lignite, peat, oil, sludge, petroleum coke, gas and wastes. All these fuels are burned cleanly and economically in CFB boilers without the need for complex scrubbers, catalytic or other costly chemical clean-up systems. Today, as CFB technology becomes more widely accepted and unit sizes increase, more focus is being placed on environmental performance of this technology. In fact, the low stack emissions are a driving force that dictates the selection of CFB boilers over pulverized coal (PC) boilers. Based upon performance data, the CFB units met and significantly exceeded all the emission levels. In most cases, operating efficiency and emission performance were better than the boiler guarantee. The development of CFB technology continues with emphasis on higher steam pressure and temperature and larger units as well as oxy-fuel combustion.

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Sources of further information and advice

Amand L -E. 1994. Nitrous oxide emission from circulating fluidized bed combustion. PhD dissertation, Chalmers University, Goeteborg, Sweden, also presented in Brereton (1996). Anthony E J, Preto F. 1995. Pressurized combustion in FBC systems, in Pressurised Fluid Bed Combustion (eds M. Alvarez Cuenca and E. J. Anthony), Glasgow: Blackie, 80–120. Dennis J S, Hayhurst A N. 1990. Mechanism of the sulphation of calcined particles in combustion gases. Chem. Eng. Sci., 45: 1175–1187. Halder P K. 1989. Combustion of single coal particles in circulating fluidized bed. PhD thesis, Technical University of Nova Scotia. Johnsson J E, Åmand L-E, Leckner B. 1990. Modelling of NO formation in a circulating fluidized bed boiler. Third Int. Conf. on Circulating Fluidized Beds, Nagoya, Japan. Oka S N. 2004. Fluidized Bed Combustion, New York: Marcel Dekker. Ściążko M., Zuwała J., Pronobis M. 2006. Zalety i wady współspalania biomasy w kotłach energetycznych na tle doświadczeń eksploatacyjnych pierwszego roku współspalania biomasy na skalę przemysłową, Energetyka i Ekologia 3: 207–220. Stantan J E. 1983. Sulfur retention in FBC, in Fluidized Beds – Combustion and Applications (ed. J R Howard), London: Applied Science.

16.8

References

Anthony E J, Preto F. 1995. Pressurized combustion in FBC systems, in Pressurised Fluid Bed Combustion (eds M. Alvarez Cuenca and E. J. Anthony), Glasgow: Blackie, 80–120. Anthony D B, Howard J B, Hottel H C, Meissner H P. 1974. Rapid devolatilization of pulverized coal. 15th Symposium (International) on Combustion, Pittsburgh: Combustion Institute, 1303–1315. Arena U, D’Amore M, Massimilla L. 1983. Carbon attrition during the fluidized combustion of a coal. AIChE Journal 29: 40–49. Avedesian M M, Davidson J F. 1973. Combustion of carbon particles in a fluidized bed. Trans. Inst. Chem. Eng. 2: 121–131. Basu P. 2006. Combustion and Gasification in Fluidized Beds, New York: Taylor and Francis Group. Basu P, Broughton J, Elliott D E. 1975. Combustion of single coal particles in fluidized beds. In: Institute of Fuel Symposium Series no. 1: Fluidised Combustion, London: Institute of Fuel, A3-1–A3-10. Bellgardt D, Hembach F, Schössler M, Werther J. 1987. Modeling of large-scale atmospheric fluidized bed combustors. 9th International Conference on Fluidized Bed Combustion, Boston, MA. Brereton C. 1996. Combustion performance. In Grace J R, Knowlton T M, Avidan A A. (eds) Circulating Fluidized Beds, London: Blackie Academic and Professional. Bryngelsson M, Westermark M. 2009. CO2 capture pilot test at a pressurized coal fired CHP plant. Energy Procedia 1: 1403–1410. © Woodhead Publishing Limited, 2013

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Campbell E K, Davidson J F. 1975. The combustion of coal in fluidized beds. In Institute of Fuel, Symposium Series No. 1, Fluidized Combustion, London: Institute of Fuel, A2-1–A2-9. Chen T P, Saxena S C. 1978. A mechanistic model applicable to coal combustion in fluidized beds. AIChE Symposium Series 176: 149–161. Czakiert T, Bis Z, Muskala W, Nowak W. 2006. Fuel conversion from oxy-fuel combustion in a circulating fluidized bed. Fuel Processing Technology 87: 531–538. Czakiert T, Sztekler K, Karski S, Markiewicz D, Nowak W. 2010. Oxy-fuel circulating fluidized bed combustion in a small pilot-scale test rig. Fuel Processing Technology 91: 1617–1623. Davidson J F, Clift R, Harrison D. 1985. Fluidization, 2nd edn. London: Academic Press. Dennis J S, Hayhurst A N. 1990. Mechanism of the sulphation of calcined particles in combustion gases. Chem. Eng. Sci., 45: 1175–1187. Diego de L F, Londono C A, Wang X S, Gibbs B M. 1996. Influence of operating parameters on NOx and N2O axial profiles in a circulating fluidized bed combustor. Fuel 75(8): 971–978. Eddings E G, Okerlund R, Bool L E. 2009. Pilot-scale evaluation of oxycoal firing in circulating-fluidized-bed and pulverized-coal-fired test facilities. Proc. 1st IEA GHG International Conference on Oxyfuel Combustion, Cottbus, Germany, September 8–11. Field M A, Gill D W, Morgan B B, Hawksley P G W. 1967. Combustion of Pulverized Coal. Leatherhead: British Coal Utilization Research Association. Glinicki M. 2001. The influence of addition of activated fly ash from CFB combustion process on the properties of structural concretes. Proc. of 8th Int. Conf. ‘Power Production Ashes’, Miedzyzdroje, Poland. Goidich S J, Fan Z, Sippu O, Bose A C. 2000. Integration of the BENSON Vertical OTU Technology and the Compact CFB Boiler. POWER-GEN International 2000, Orlando, FL, November 14–16. Hilger A. 1991. FBC Technology and the Environmental Challenge. Proceedings of the Institute of Energy’s Fifth International Fluidized Combustion Conference, London, 10–11 December 1991, p. 383. Hoffman A C, Stein L E. 2002. Gas Cyclones and Swirl Tubes. New York: Springer. Horio M, Wen C Y. 1978. Simulation of fluidized bed combustors: Part 1. Combustion efficiency and temperature profile. AIChE Symposium Series 176: 101–111. Hotta A. 2009. Foster Wheeler’s solution for large-scale CFB boiler technology: features and operational performance of Lagisza 460MWe CFB boiler. The 20th International Conference on Fluidized Bed Combustion, Xi’an China: 59–70. Johnsson J E, Åmand L-E, Leckner B. 1990. Modelling of NO formation in a circulating fluidized bed boiler. Third Int. Conf. on Circulating Fluidized Beds, Nagoya, Japan. Kledyński Z, Ziarkowska K. 2002. 1st Congress of Environmental Engineering, Lublin. Kobyłecki R, Nowak W, Maslanka J, Ziarkowska K. 2004. Polish experiences in utilization of ashes from fluidized bed coal combustion. Proc. of Colloquium ‘Fluidized Bed Combustion and Gasification for Power Generation’, South Africa. Krzywanski J, Czakiert T, Muskala W, Sekret R, Nowak W. 2010. Modeling of solid fuels combustion in oxygen-enriched atmosphere in circulating fluidized bed boiler – Part I. The mathematical model of fuel combustion in oxygen-enriched CFB environment. Fuel Processing Technology 91: 290–295. Lapple C E. 1951. Process use many collector types. Chem. Eng., 58: 144–151.

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Majchrzak-Kuceba I. 2001. Experimental investigations of removal and disposal of CO2 from flue gases using zeolites, PhD thesis, Czestochowa University of Technology, Poland. Mirek P, Sekret R, Nowak W. 2007. The influence of air nozzles’ shape on the NOx emission in the large-scale 670 MWth CFB boiler. Fluidization XII, New Horizons in Fluidization Engineering: 969–976. Muschelknautz E, Trefz M. 1993. Pressure drop and separation efficiency in cyclones. VDI Heat Atlas, 6th edn, Dusseldorf: VDI Verlag. Muschelknautz E, Greif V, Trefz M. 1994. Zyklone zur Abscheidung von Feststoffen aus Gasen. VDI Heat Atlas, 7th edn, Dusseldorf: VDI-Verlag. Nowak W, Bis Z. 1997. Analysis of Operating Conditions of OFz-450 CFB Boiler, Report No. IT/2985/97, T.U. of Czestochowa (in Polish). NUCLA Circulating Fluidized-Bed Demonstration Project, A US DOE Post-Project Assessment, June 1995, DOE/METC-95/1019 (DE95009706). NUCLA ACFB Demonstration Project, Project Performance Summary, Clean Coal Demonstration Program, June 1999, DOE/FE-0392. Oka S N. 2004. Fluidized Bed Combustion, New York: Marcel Dekker. Rajczyk R., Bednarek M. 2011. Spalanie i współspalanie biomasy stałej w energetyce. Energia i Budynek (9), 3–8. Report of EVT No. V4945, GEC Alstrom, 1997. Report of the Polish Institute of Energy, No. 89 374 95C/2388, 1997 (in Polish). Ross I B, Davidson J F. 1981. The combustion of carbon particles in a fluidized bed. Trans. Inst. Chem. Eng 59: 108–114. Skowyra R S, Tanca M C. 1993. Circulating fluidized bed steam generator experience overview, ABB Technical paper TIS 93-S1. Stairmand C J. 1951. The design and performance of cyclone separators. Trans. Instn. Chem. Engrs. 29: 356–383. Stubington J F, Clough S J. 1997. The combustion rate of volatiles in a fluidized bed combustor. Proceedings of the 14th International Conference on Fluidized Bed Combustion, F. D. S. Preto (ed), 2: 1111–1122. Suzuki Y, Moritomi H, Kido N. 1991. On the formation mechanism of N 2O during circulating fluidized bed combustion. Proceedings of 4th SCEJ Symposium on CFB, Japan. Swift P. 1969. Dust control in industry. Steam Heat Engineer 38: 453–456. Szymanek A. 2002. Sorbents for power sector. Proc. of Ecoenergia Fairs, Gdansk, Poland. Tan Y, Jia L, Wu Y, Anthony E J. 2012. Experiences and results on a 0.8 MWth oxy-fuel operation pilot-scale circulating fluidized bed. Applied Energy, 92: 343–347. Turnbull E, Kossakowski E R, Davidson J F, Hopes R B, Blackshow H W, Goodyear P T Y. 1984. Effect of pressure on combustion of char in fluidized beds. Chem. Eng. Research & Development 7: 223–233. Zheng L. (ed.) 2011. Oxy-fuel Combustion for Power Generation and Carbon Dioxide (CO2) Capture. Cambridge: Woodhead Publishing.

16.9

Appendix: nomenclature

16.9.1 Symbols as

emissivity of gas-solids phase, dimensionless

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Circulating fluidized bed combustion (CFBC)

aw A Afree ASH ASHn Aw AF cm cp,g cp,s d50 Ddiff de de* Dfr dfr di* dk E E0 Ec Esor Fc Fsor Fu FOF h Hf upper H cir lower H cir

i j k katt Katt kb kc ke kf ko

761

emissivity of heat exchanger surface, dimensionless boiler cross section area, m2 surface area of free space inside combustion chamber, m2 superheater surface area, m2 surface area of walls combined to the superheater, m2 overall surface area of walls in combustion chamber, m2 availability factor specific heat capacity, J/kg K specific heat capacity of gas, J/kg◊k specific heat capacity of solids, J/kg◊k mean particle size of the feed, m oxygen diffusion coefficient, m2/s equivalent diameter of combustion chamber, m cut size for wall separation, m fragmentation factor, dimensionless specific diameter, m cut size for separation in the vortex, m the diameter of the k-th class of particles, m activation energy, kJ/mol average activation energy, kJ/mol average cyclone efficiency sulfur capture efficiency coal feed rate, kg/sec sorbent feed rate, kg/sec fouling index forced outage factor overall heat-transfer coefficient, W◊m–2◊k–1 height of the furnace above the secondary air level, m solid flow enthalpy in the upper part of the combustion chamber, kJ/kg solid flow enthalpy in the lower part of the combustion chamber, kJ/kg i-th independent reaction for i-th component of volatiles, which has specific activation energy, i.e., reaction rate index of particle class after fragmentation, dimensionless exponent for limiting loading ratio mlim,dimensionless attrition rate, m/s attrition constant, 1/m bed material circulation ratio, dimensionless combustion rate constant for char, s/m elutriation rate constant, m/s mass transfer coefficient, s/m pre-exponential factor for char combustion, s/m © Woodhead Publishing Limited, 2013

762

kRi kRo l LHV m· mbcomp m·cir m· f mg ms M n1 n2 nAi Ofurn P Pfr(dk) pg Q0cir QIcir I QIcir out Q cir QIf Qinf Qfg QIfg Qfgout Qa Q0a QIa QIIa Qair QE Qb Qba Qbott Qreact

Fluidized bed technologies for near-zero emission combustion

rate of reaction pre-exponential coefficient, s–1 height of the membrane-wall, m lower heating value, J/kg mass flow rate of a specified solid fraction, kg/s mass of component in the bed, kg mass flow of solids circulating in the CFB loop, kg/s mass flow of fuel introduced into the furnace, kg/s mass of gas, kg mass of solids, kg molar mass, g/mol primary fragmentation coefficient, dimensionless secondary fragmentation coefficient, dimensionless exponent in RRSB distribution of feed, dimensionless combustion chamber perimeter, m proportionality constant in pore plugging time s (=P ¥ (SO2 concentration, kmol/m3)–1) probability of fragmentation of the kth class, dimensionless oxygen partial pressure, Pa energy of circulating solids flowing into lower part of combustion chamber, kW energy of circulating solids flowing into combustion chamber, kW energy of circulating solids flowing into a cyclone separator, kW energy of circulating solids flowing out of the convection section, kW energy of fuel flowing into lower part of combustion chamber, kW energy of fuel flowing into combustion chamber, kW energy of flue gas flowing into the convective section, kW energy of flue gas flowing into combustion chamber, kW energy of flue gas flowing out of the convective section, kW energy of air flowing out of the LUVO, kW energy of air flowing into the LUVO, kW energy of primary air flowing into lower part of combustion chamber, kW energy of secondary air flowing into combustion chamber, kW energy of air, W energy of heat exchanger in combustion chamber, kW energy of heat exchanger in the convective section, kW energy of the bottom ash flowing out of the bottom part of the combustion chamber, kW energy of bottom ash, W energy of reactions, W

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Circulating fluidized bed combustion (CFBC)

Qrec Qv Qwall Pr RR RAi(d) Rg s s SOF t ta tb tk tp tr ts T Tupper Tlower Tp Tw Ts Tg Tgs Twz U Ug Us Ut Vfree Vfurn Vg VM VM* w wfr XcaCO3 z a1 aconv ag ap

763

energy of recycled particles, W energy used for moisture vaporization, W energy lost through walls, W Prandtl number for flue gas, – recirculation ratio, dimensionless residue of the ‘inner feed’, dimensionless universal gas constant, 8.314 J/mol k thickness of radiating layer, m sulfur fraction in coal serviced outage factor time, s forced outage time, h actual service outage time, h length of calculation period, h run time, h planned production limitations, h average service outage, h temperature, k temperature in the upper part of the combustion chamber, k temperature in the lower part of the combustion chamber, k fluidized bed temperature, K membrane-wall temperature, k temperature of the coarse fuel particle, k average flue gas temperature, K gas temperature in surface boundary layer, k temperature of the dirty wall, k superficial gas velocity, m/s flue gas velocity, m/s particle velocity, m/s terminal velocity, m/s volume of free space inside combustion chamber, m3 combustion chamber volume, m3 real volume of a flue gas for the combustion air factor, m3/kg amount of volatile yet to be released, % total volatiles content, % average flue gas velocity, m◊s–1 fragmentation coefficient, dimensionless weight fraction of calcium carbonate in the sorbent particle distance of cross-section area from the grid, m heat-transfer coefficient in the dense zone, W◊m–2◊k–1 heat-transfer coefficient due to convection, W◊m–2◊k–1 heat-transfer coefficient of flue gas, W◊m–2◊k–1 heat-transfer coefficient of bed material particles, W◊m–2◊k–1

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764

arad dmax ef hcyc hi hlim htot lI lc me mlim r* rbav rc rg rp sG tg tp n d l x t y

Fluidized bed technologies for near-zero emission combustion

heat-transfer coefficient due to radiation, W◊m–2◊K–1 maximum extent of sulfation fouling coefficient, m2◊K◊W–1 cyclone efficiency, dimensionless separation efficiency due to inner vortex, dimensionless separation efficiency due to wall, dimensionless total separation efficiency, dimensionless primary excess air thermal conductivity for flue gas, W◊m–1◊K–1 loading ratio in cyclone entrance, dimensionless limited loading ratio, dimensionless mass concentration of ash connected with the solids circulation rate, kg/m3 average bed density, kg/m3 coal density, kg/m3 flue gas density, kg/m3 density of bed material, kg/m3 standard deviation of Gaussian distribution of activation energy, kJ/mol optical thickness of flue gas, dimensionless optical thickness of gas-solids phase, dimensionless coefficient of kinematic viscosity for flue gas, m2◊s–1 thickness of surface boundary layer, m excess air coefficient of exploitation considering non-uniform flow of flue gas in a short distance to the heat exchanger surface, dimensionless optical thickness, dimensionless coefficient considering the internal circulation of bed material in the combustion chamber, dimensionless

16.9.2 Subscripts ash att c com D e in R RD RFA s

ash attrition coal combustion drainage elutriation input flow return through the return leg recirculation in drainage recirculation in fly ash sorbent

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