A typical gasification process generally follows the sequence of steps listed on the next page (illustrated schematically in Figure 5.2).

• Preheating and drying

• Pyrolysis

• Char gasification

• Combustion

Though these steps are frequently modeled in series, there is no sharp boundary between them, and they often overlap. The following paragraphs discuss these sequential phases of biomass gasification.

In a typical process, biomass is first heated (dried) and then it undergoes thermal degradation or pyrolysis. The products of pyrolysis (i. e., gas, solid, and liquid) react among themselves as well as with the gasifying medium to form the final gasification product. In most commercial gasifiers, the thermal energy necessary for drying, pyrolysis, and endothermic reactions comes from a certain amount of exothermic combustion reactions allowed in the gasifier. Table 5.2 lists some of the important chemical reactions taking place in a gasifier.

The bubbling fluidized-bed gasifier, developed by Fritz Winkler in 1921, is perhaps the oldest commercial application of fluidized beds; it has been in commercial use for many years for the gasification of coal (Figure 6.8); for biomass gasification, it is one of the most popular options. A fairly large number of bubbling fluidized-bed gasifiers of varying designs have been developed and are in operation (Lim and Alimuddin, 2008; Narvaez et al., 1996).

Because they are particularly suitable for medium-size units (< 25 MWth), many biomass gasifiers operate on the bubbling fluidized-bed regime. Depend­ing on operating conditions, bubbling-bed gasifiers can be grouped as low — temperature and high-temperature types. They can also operate at atmospheric or elevated pressures.

In the most common type of fluidized bed, biomass crushed to less than 10 mm is fed into a bed of hot materials. These bed materials are fluidized with steam, air, or oxygen, or their combination, depending on the choice of

gasification medium. The ash generated from either the fuel or the inorganic materials associated with it is drained easily from the bottom of the bed. The bed temperature is normally kept below 980 °C for coal and below 900 °C for biomass to avoid ash fusion and consequent agglomeration.

The gasifying medium may be supplied in two stages. The first-stage supply is adequate to maintain the fluidized bed at the desired temperature; the second — stage supply, added above the bed, converts entrained unreacted char particles and hydrocarbons into useful gas.

High-temperature Winkler (HTW) gasification is an example of high — temperature, high-pressure bubbling fluidized-bed gasification for coal and lignite. Developed by Rheinbraun AG of Germany, the process employs a pres­surized fluidized bed operating below the ash-melting point. To improve carbon conversion efficiency, small char particles in the raw gas are separated by a cyclone and returned to the bottom of the main reactor (Figure 6.9).

The gasifying medium (steam and oxygen) is introduced into the fluidized bed at different levels as well as above it. The bed is maintained at a pressure of 10 bars while its temperature is maintained at about 800 °C to avoid ash fusion. The overbed supply of the gasifying medium raises the local temperature to about 1000 °C to minimize production of methane and other hydrocarbons.

The HTW process produces a better-quality gas compared with the gas that is produced by traditional low-temperature fluidized beds. Though originally

developed for coal, it is suitable for lignite and other reactive fuels like biomass and treated municipal solid waste (MSW).

Chemical reactions involve the mixing of reactants. If the mixing is incomplete, the reaction will be incomplete, even if the right amounts of reactant and the right temperature are available. The mixing is better when all reactants are either in the gas phase or in the liquid phase compared to that when one reactant is in the solid phase and the other is in the gas or liquid phase. The absence of interphase resistance in a monophase reaction medium greatly improves the mixing. The conventional thermal gasification of solid biomass in air or steam involves heterogeneous mixing, and therefore the gas-solid interphase resis­tance limits the conversion reactions.

Supercritical water allows reactions to take place in a single phase, as most organic compounds and gases are completely miscible in it. It is thus a superior reaction medium. Because the absence of interphase mass transfer resistance facilitates better mixing and therefore higher conversion, SCW can be an excel­lent medium for the following three types of reactions:

Hydrothermal gasification of biomass. SCW is an ideal medium for gas­ification of very wet biomass, such as aquatic species and raw sewage, which ordinarily have to be dried before they can be gasified economically. SCW gasification produces gas at high pressure and thus obviates the need for an expensive product gas compression step for transport or use in combustion.

Synthesis reactions. A variety of organic reactions like hydrolysis and molecular rearrangement can be effectively carried out in SCW, which serves as a solvent, a reactant, and sometimes a catalyst. There is no need for acid or base solvents, the disposal of which is often a problem. Supercritical water oxidation. Complete miscibility of oxygen in SCW helps harmful organic compounds to be easily oxidized and degraded. Thus, SCW is an attractive means of turning pollutants into harmless oxides.

To achieve mass flow, the following conditions are to be met:

• The hopper wall must be sufficiently smooth for mass flow.

• The hopper angle should be adequately steep to force solids to flow at the

walls.

• The hopper outlet must be large enough to prevent arching.

• The hopper outlet must be adequately large to achieve the maximum dis­

charge rate.

The required smoothness and sloping angle for mass flow in a hopper depends on the friction between the particles and the hopper surface. This friction can be measured in a laboratory using a standard test (ASTM, 2000).

Several factors can affect wall friction for a given fuel:

• Wall material

• Surface texture or roughness of the wall

• Moisture content and variations in solids composition and particle size

• Length of time solids remain unmoved

• Corrosion of wall material due to reaction with solids

• Scratching of wall material caused by abrasive materials

To enhance the smoothness of the surface, sometimes the hopper is coated or a smooth lining is applied. Lining materials that can be used include polyure­thane sheets, TIVAR-88, ultra-high-molecular-weight polyethylene plastic, and krypton polyurethane.

Mass flow can be adversely affected by the narrowness of the hopper outlet. A too-narrow outlet (compared to particle size) permits the particles to interlock when exiting and form an arch over the outlet. The probability of this happening increases when

• The particles are large compared to the outlet width.

• There is high moisture in the solids.

• The particles are of a low shape factor and have a rough surface texture.

• The particles are cohesive.

Wedge-shaped hoppers require a smaller width than conical hoppers do in order to prevent bridging. Slotted outlets must be at least three times as long as they are wide.

Hopper Design for Mass Flow

The design of the hopper outlet significantly affects the flow of solids. When solids flow through the hopper, air or gas enters, dilating the particles. It is essential for powder solids to flow freely through the outlet. Air drag, which is proportional to surface area, must be balanced by gravitational force that is equal to the weight of the particle. Fine particles have a lower ratio of weight to surface area compared to coarser particles. So, for fine particles, this force balance becomes an important consideration. For such particles, the following expression is used (Carleton, 1972):

where

Vo is the average solid velocity through the outlet, m/s

pa, pp is the density of the air and solids, respectively, kg/m3

dp is the particle size, m

Ц is the viscosity of the air, kg/m. s

в is the semi-included angle of the hopper

g is the acceleration due to gravity, 9.81 m/s2

B is the parameter

The mass-flow rate, m, is given in terms of the bulk solid density, pb, and the outlet area, A:

m = pbAVo (8.2)

For coarse particles (>500 |jm), an alternative relation is used (Johanson, 1965):

Design Steps

Hopper design involves determining particle properties, such as interparticle friction, particle-to-wall friction, and particle compressibility or permeability. With these properties known, the outlet size, hopper angle, and discharge rate are found.

Dedicated experiments like shear tests are carried out to determine the inter­particle friction. A parameter, such as angle of repose, has little value in hopper

design, as it simply gives the heap angle when solids are poured in. Particle-wall friction should also be measured by purpose-designed experiments.

The stress distribution on the silo wall is important, especially for a tall unit. Figure 8.7 compares the wall pressure in a biomass-filled silo with that of a liquid-filled silo. As we can see, the wall pressure in a solid-filled silo does not vary linearly with height, but it does in a liquid-filled silo. In the former case, the pressure increases with depth, reaching an asymptotic value that depends on the diameter of the hopper rather than on the height. Because there is no further increase in wall stress with height, large silos are smaller in diameter but taller.

To find the stress in the barrel, or the vertical wall section, of a hopper, we consider the equilibrium of forces on a differential element, dh, in a straight­sided silo (Figure 8.8):

• Vertical force due to pressure acting from above: Pv A

• Weight of material in element: pAg dh

• Vertical force due to pressure acting from below: (Pv + dPv) A

• Solid friction on the wall acting upward: mD dh

The force balance on the elemental solid cross-section gives

(Pv + dPv ) A + mD dh = PvA + pAg dh (8.4)

The wall friction is equal to the particle-wall friction coefficient, kf, times the normal pressure on wall, Pw:

T = kfPw (8.5)

Janssen (1895) assumed the lateral pressure to be proportional to the vertical pressure, as shown in the following equation:

FIGURE 8.8 Force balance on an element of a storage silo.

Pw = kpv

where K is the Janssen coefficient. For liquids, the pressure is uniform in all directions, so K is 1.0. This relation is not strictly valid for all solids, but for engineering approximations we can start with this assumption.

Substituting Eqs. (8.5) and (8.6) in Eq. (8.4), we get

AdPv = pAg dh — kfKPvnD dh (8.7)

Boundary conditions for this equation are h = 0, Pv = 0; h = H, Py = P0. With this, Eq. (8.7) is integrated from h = 0 to h = H to get the pressure at the base of the silo’s vertical section, P0. Substituting

we get

This is known as the Janssen equation.

Figure 8.7 illustrated the pressure distribution along the height of a silo. The straight line shows the pressure we expect if the stored substance is a liquid; the discontinuous exponential curve is the one predicted for solids. There is a sharp increase in pressure at the beginning of the inclined wall. The pressure decreases with height (Figure 8.7).

The stress on the inclined section is different from that calculated from the preceding. To calculate this, we use the Jenike equation, which states that the radial pressure is proportional to the distance of the element from the hopper apex, which is the point where inclined surfaces would meet if they were

extended (Jenike, 1964). It can be seen that the magnitude of stress at the hopper exit is the lowest, although this is the lowest point in the hopper.

Example 8.1

Find the wall stress at the bottom of a large silo, 4.0 m in diameter and 20 m in height, that uses a flat bottom for its discharge. Compare the stress when the silo is filled with wood chips (bulk density 300 kg/m3) with that when it is filled with water.

Given that the wall-to-wood chip friction coefficient, kf, is 0.37, assume the Janssen coefficient, K, to be 0.4.

Solution

We use Eq. (8.8) to calculate the vertical pressure, P0, in the silo. Data given are as follows:

* The bulk density of the wood chips, p, is 300 kg/m3.

* The wall-solid friction coefficient, kf, is 0.37.

* The diameter, D, is 4.0 m.

* The height, H, is 20 m.

*

The Janssen coefficient, K, is 0.4.

Since the lateral pressure, Pw, is proportional to the vertical pressure, Pv,

Pw = KP0 = 0.4 x 1 8,854 = 7542 Pa For water, the vertical pressure is the weight of the liquid column:

P0 =PgH

Because the lateral and vertical pressures are the same (i. e., K = 1.0), we can write

Pw = P0 = 1000 X 9.81 x 20 = 196,200 Pa

The lateral pressure for water is therefore (196,200/7542) or 26 times greater than that for wood chips.

Methanol may be converted into gasoline using several processes. One of these, Exxon Mobil’s methanol-to-gasoline (MTG) process, is well known (Figure 9.4). Methanol is converted into hydrocarbons consisting of mainly (>75%) gasoline-grade materials (C5-Q2) with a small amount of liquefied petroleum gas (C3-C4) and fuel gas (Cj-C2). Mobil uses both fixed beds and fluidized beds

of proprietary catalysts for this conversion. The reaction is carried out in two stages: the first stage is dehydration to produce dimethyl ether intermediate; the second stage is also dehydration, this time over a zeolite catalyst, ZSM-5, to give gasoline.

where (-H2O) represents the dehydration step.

The typical composition of the gasoline in weight percentage (see nzic. org. nz/ChemProcesses/energy/7D. pdf) is as follows:

• Highly branched alkanes: 53%

• Highly branched alkenes: 12%

• Napthenes: 7%

• Aromatics: 28%

The dehydration process produces a large amount of water. For example, from 1000 kg of methanol, 387 kg of gasoline, 46 kg of liquefied petroleum gas, 7 kg of fuel gas, and 560 kg of water are produced (Adrian et al., 2007). Figure 9.4 shows a simplified scheme for the production of gasoline from methanol. This gasoline, sometimes referred to as MTG gasoline, is completely compatible with petrogasoline.

Gasifier simulation models may be classified into the following groups:

• Thermodynamic equilibrium

• Kinetic

• Computational fluid dynamics (CFD)

• Artificial neural network

The thermodynamic equilibrium model predicts the maximum achievable yield of a desired product from a reacting system (Li et al., 2001). In other words, if the reactants are left to react for an infinite time, they will reach equilibrium yield. The yield and composition of the product at this condition is given by the equilibrium model, which concerns the reaction alone without taking into account the geometry of the gasifier.

In practice, only a finite time is available for the reactant to react in the gasifier. So, the equilibrium model may give an ideal yield. For practical appli­cations, we need to use the kinetic model to predict the product from a gasifier that provides a certain time for reaction. A kinetic model studies the progress of reactions in the reactor, giving the product compositions at different posi­tions along the gasifier. It takes into account the reactor’s geometry as well as its hydrodynamics.

CFD models (Euler type) solve a set of simultaneous equations for conser­vation of mass, momentum, energy, and species over a discrete region of the gasifier. Thus, they give distribution of temperature, concentration, and other parameters within the reactor. If the reactor hydrodynamics is well known, a CFD model provides a very accurate prediction of temperature and gas yield around the reactor.

Neural network analysis is a relatively new simulation tool for modeling a gasifier. It works somewhat like an experienced operator, who uses his or her years of experience to predict how the gasifier will behave under a certain condition. This approach requires little prior knowledge about the process. Instead, the neural network learns by itself from sample experimental data (Guo et al., 1997).

Thermodynamic Equilibrium Models

Thermodynamic equilibrium calculation is independent of gasifier design and so is convenient for studying the influence of fuel and process parameters. Though chemical or thermodynamic equilibrium may not be reached within the gasifier, this model provides the designer with a reasonable prediction of the maximum achievable yield of a desired product. However, it cannot predict the influence of hydrodynamic or geometric parameters, like fluidizing velocity, or design variables, like gasifier height.

Chemical equilibrium is determined by either of the following:

• The equilibrium constant

• Minimization of the Gibbs free energy

Prior to 1958 all equilibrium computations were carried out using the equilib­rium constant formulation of the governing equations (Zeleznik and Gordon,

1968). Later, computation of equilibrium compositions by Gibbs free energy minimization became an accepted alternative.

This section presents a simplified approach to equilibrium modeling of a gasifier based on the following overall gasification reactions:

R1: CO2 + C ^ 2 CO (5.61)

R2:C + H2O ^ H2 + CO (5.62)

R3:C + 2H2 ^ CH4 (5.63)

R9:CO + H2O ^ CO2 + H2 (5.64)

From a thermodynamic point of view, the equilibrium state gives the maximum conversion for a given reaction condition. The reaction is considered to be zero dimensional and there are no changes with time (Li et al., 2001). An equilibrium model is effective at higher temperatures (> 1500 K), where it can show useful trends in operating parameter variations (Altafini et al., 2003). For equilibrium modeling, one may use stoichiometric or nonstoichiometric methods (Basu,

2006) .

Stoichiometric Equilibrium Models

In the stoichiometric method, the model incorporates the chemical reactions and species involved. It usually starts by selecting all species containing C, H, and O, or any other dominant elements. If other elements form a minor part of the product gas, they are often neglected.

Let us take the example of 1 mole of biomass being gasified in d moles of steam and e moles of air. The reaction of the biomass with air (3.76 moles of nitrogen, 1 mole of oxygen) and steam may then be represented by

CHaO„Nc + dH2O + e (O2 + 3.76N2 )^ n1C + n2H2 + n3CO 5 65)

+ N4H2O + N5CO2 + n6CH4 + N7N2

where П1…П7 are stoichiometric coefficients. Here, CHaOsNc is the chemical representation of the biomass and a, b, and c are the mole ratios (H/C, O/C, and N/C) determined from the ultimate analysis of the biomass. With d and e as input parameters, the total number of unknowns is seven.

An atomic balance of carbon, hydrogen, oxygen, and nitrogen gives

During the gasification process, reactions R1, R2, R3, and R9 (see Table 5.2) take place. The water-gas shift reaction, R9, can be considered a result of the

subtraction of the steam gasification and Boudouard reactions, so we consider the equilibrium of reactions R1, R2, and R3 alone. For a gasifier pressure, P, the equilibrium constants for reactions Rb R2, and R3 are given by

where yi is the mole fraction for species i of CO, H2, H2O, and CO2.

The two sets of equations (stoichiometric and equilibrium) may be solved simultaneously to find the coefficients, (nj…n7), and hence the product gas composition in an equilibrium state. Thus, by solving seven equations (Eqs. 5.66-5.72) we can find seven unknowns (nj…n7), which give both the yield and the product of the gasification for a given air/steam-to-biomass ratio. The approach is based on the simplified reaction path and the chemical formula of the biomass.

This is a greatly simplified example of the stoichiometric modeling of a gasification reaction. The complexity increases with the number of equations considered. For a known reaction mechanism, the stoichiometric equilibrium model predicts the maximum achievable yield of a desired product or the pos­sible limiting behavior of a reacting system.

The amount of gasification medium has a major influence on yield and com­position of the product gas. This section discusses methods for choosing that amount.

Air

The theoretical air requirement for complete combustion of a unit mass of a fuel, mth, is an important parameter. It is known as the stoichiometric air requirement. Its calculation is shown in Eq. 2.34. For an air-blown gasifier operating, the amount of air required, Ma, for gasification of unit mass of biomass is found by multiplying it by another parameter ER:

Ma = mthER (6.10)

Here, ER is the equivalence ratio.

For a fuel feed rate of Mf, the air requirement of the gasifier, Mfa, is

Mfa = mthER ■ Mf (6.11)

For a biomass gasifier, 0.25 may be taken as a first-guess value for the equiva­lence ratio, ER. A more detailed discussion of this is presented next.

Equivalence Ratio The equivalence ratio is an important gasifier design parameter. It is the ratio of the actual air-fuel ratio to the stoichiometric air-fuel ratio. This term is generally used for air-deficient situations, such as those found in a gasifier.

where EA is the excess air coefficient.

In a combustor, the amount of air supplied is determined by the stoichio­metric (or theoretical) amount of air and its excess air coefficient. In a gasifier, the air supply is only a fraction of the stoichiometric amount. The stoichiometric amount of air may be calculated based on the ultimate analysis of the fuel.

The equivalence ratio, ER, dictates the performance of the gasifier. For example, pyrolysis takes place in the absence of air and hence the ER is zero; for gasification of biomass, it lies between 0.2 and 0.3.

Downdraft gasifiers give the best yield for ER, 0.25 (Reed and Das, 1988, p. 25). With a lower ER value, the char is not fully converted into gases. Some units deliberately operate with a low ER to maximize their charcoal production. A lower ER gives rise to higher tar production, however, so updraft gasifiers, which typically operate with an ER of less than 0.25, have higher tar content. With an ER above 0.25, some product gases are also burnt, increasing the temperature.

The quality of gas obtained from a gasifier strongly depends on the ER value, which must be significantly below 1.0 to ensure that the fuel is gasified rather than combusted. However, an excessively low ER value (<0.2) results in several problems, including incomplete gasification, excessive char formation, and a low heating value of the product gas. On the other hand, too high an ER (>0.4) results in excessive formation of products of complete combustion,

100!

g >- о c Ш

Ъ

Ш c о ‘со ш > о

.Q

FIGURE 6.20

such as CO2 and H2O, at the expense of desirable products, such as CO and H2. This causes a decrease in the heating value of the gas. In practical gasification systems, the ER value is normally maintained within the range of 0.20 to 0.30. Figure 6.20 shows the variation in carbon conversion efficiency of a circulating fluidized-bed gasifier for wood dust against the equivalence ratio. The efficiency increases with ER and then it starts declining. The optimum value here is 0.26, but it may change depending on many factors.

The bed temperature of a fluidized-bed gasifier increases with the ER because the higher the amount of air, the greater the extent of the combustion reaction and the higher the amount of heat released (Figure 6.21). Example 6.1 illustrates the calculation procedure for ER.

Oxygen

Oxygen is used primarily to provide the thermal energy needed for the endo­thermic gasification reactions. The bulk of this heat is generated through the following partial and/or complete oxidation reactions of carbon:

C + 0.5 O2 ^ CO -111 kJ/mol (6.13)

C + O2 ^ CO2 — 394 kJ/mol (6.14)

It can be seen that for the oxidation of 1 mol of carbon to CO2, the oxygen requirement is (2 x 16)/12 = 2.66 mol, while that for carbon to CO is (16/12) = 1.33 mol. Thus, the reaction in Eq. (6.13) is more likely to take place in oxygen-deficient regions.

Besides supplying the energy for the endothermic gasification reactions, the gasifier must provide energy to raise the feed and gasification medium to the reaction temperature, as well as to compensate for the heat lost to the reactor walls. For a self-sustained gasifier, part of the chemical energy in the biomass

Equivalence ratio

FIGURE 6.21 Gasifier temperature in a CFB riser increases with equivalence ratio.

provides the heat required. The total heat necessary comes from the oxidation reactions. The energy balance of the gasifier is thus the main consideration in determining the oxygen-to-carbon (O/C) ratio.

Equilibrium calculations can show that as the ratio of oxygen to carbon in the feed increases, CH4, CO, and hydrogen in the product decreases but CO2 and H2O in the product increases. Beyond a ratio of 1.0, hardly any CH4 is produced.

When air is the gasification medium, as is the case for 70% of all gasifiers (Ciferno and Marano, 2002), the nitrogen in it dilutes the product gas. The heating value of the gas is therefore relatively low (4-6 MJ/m3). When pure oxygen from an air-separation unit is used, the heating value is higher, in the range 10 to 15 MJ/m3, but a large amount of energy (~2.18 MJ/kg O2) is spent in separating the oxygen from the air (Grezin and Zakharov, 1988).

The oxygen requirement of a gasifier can be met by either air supply or an air-separation unit that extracts oxygen from air.

Steam

Superheated steam as a gasification medium is used either alone, with air, or with oxygen. It contributes to the generation of hydrogen.

The quantity of steam, Mfh, is known from the steam-to-carbon (S/C) molar ratio.

where Mf is the fuel feed rate, and C is the carbon fraction in the fuel.

The S/C mole ratio has an important influence on product composition, as the ER has. Both hydrogen and CO increase with an increasing S/C ratio for a given temperature and oxygen-to-carbon molar ratio. The production of these two gases increases with decreasing pressure, decreasing oxygen, and decreas­ing S/C ratio. However, there is only a marginal gain in increasing the S/C molar ratio above 2 to 3, as the excess steam simply leaves the gasifier unre­acted (Probstein and Hicks, 2006, p. 119). So a value in the range of 2.0 to 2.5 can give a reasonable starting value. [4] [5]

S/C molar ratio = 107.22/48.31 = 2.22 O/C molar ratio = 12.5/48.31 = 0.26

To find the stoichiometric oxygen requirement, the oxygen required to oxidize carbon to CO2, hydrogen to H2O, and sulfur to SO2 has be to calculated.

* Twelve kg of carbon (1 mol) react with 32 kg of oxygen (1 mol) to produce 1 mol of CO2:

c + o2 = co2

Therefore, the oxygen required for 1 kg of carbon is 32/12.

* Thirty-two kg of sulfur (1 mol) react with 32 kg of oxygen (1 mol) to produce 1 mol of SO2:

s + o2 = so2

Therefore, the oxygen required for 1 kg of sulfur is 32/32 = 1.

* Similarly, 4 kg of hydrogen react with 32 kg of oxygen to produce H2O:

2H2 + 02 = 2H20

Therefore, the oxygen required for 1 kg of hydrogen is 32/4 = 8.

+ 8x 0.059 + 0.043- 0.111 = 2.465 kg of 02/kg of fuel The total O2 required is

750 x 2.465 = 1 848.75 kg of 02/min

The O2 supplied is

moles of 02 x 32 = 12.5 x 32 = 400 kg of 02/min From this we can calculate

ER = 400/1848.75 = 0.22

The syngas constituents in the total product gas are CO (15.2%) and H2 (42.3%). So, to produce 1000 Nm3/min of syngas, the amount of product gas,

Qpr, is

Qpr = 1 000/(0.1 52 + 0.423) = 1739 Nm3/min The cross-sectional area of the gasifier reactor, A, is

A = n 474 = 12.56 m2

Assuming the operating temperature to be 1000 °C and the pressure to be 25 bars, the volumetric flow rate of product gas is

The space velocity of the gas flow Vg is Qpr/ A = 324/ (12.56 x 60) = 0.43 m/s.

The energy produced per Nm3 of product gas is found by multiplying the volume fraction by the heating value of each constituent, which is taken from Table C.2 in Appendix C. Adding together the contribution of all product gas constituents gives the total heating value, HHV, as

HHV = 0.004 x 25.1 + 0.152 x(282.99/22.4) + 0.423 x(285.84/22.4)

+ 0.086x(890.36/22.4) + 0.008×63.4 = 11.33 MJ/Nm3

Thus, the total energy produced, £№/, is Qpr x HHV

= 1739 x 11.33/60 = 328.3 MWth

The hearth load is

Etotal/A = 328.3/12.56 = 26.14 M^m2

Limited data obtained by Sinag et al. (2004) suggest that at a higher heating rate the yield of hydrogen, methane, and carbon dioxide increases while that of carbon monoxide decreases. Further investigation is needed to elucidate this point.

7.4.2 Feed Particle Size

The effect of biomass particle size is not well researched. With limited data, Lu et al. (2006) showed that smaller particles result in a slightly improved hydrogen yield and higher gasification efficiency. However, Mettanant et al. (2009) did not observe any effect when they varied the size of rice husk par­ticles in the range of 1.25 to 0.5 mm. Even if the size effect is confirmed with further data, it remains to be seen if the extra energy required for grinding is worth the improvement.

г ; л

TABLE 7.3 Effect of Solid Content in Feed and Other Operating Parameters on Gasification

7.4.3 Pressure

Experiments by Van Swaaij et al. (2003) in their microreactor over the range of 19 to 54 MPa, those by Kruse et al. (2003) in a stirred tank (30-50 MPa, 500 °C), and those by Lu et al. (2006) in a plug-flow reactor (18-30 MPa, 625 °C) showed no major effect of pressure on carbon conversion or product distribution. Nor did Mettanant et al. (2009) see much effect in their tempera­ture and pressure range, although they noted a clear positive effect of pressure at 700 °C. This issue needs further exploration.

The overbed system (Figure 8.20a) is simple, reliable, and economical, but it causes a loss of fine biomass particles through entrainment. In this system, the top size of the fuel particles is coarser than that in an underbed system, making

FIGURE 8.20 Position of feeders in a bubbling fluidized bed: (a) overbed feeding and (b) under­bed feeding.

fuel preparation simpler and less expensive. However, the feed can contain a large amount of fines with a terminal velocity that is higher than the superficial velocity in the freeboard. When the terminal velocity is lower than the super­ficial velocity of the fluidized bed, the particles are elutriated before they completely gasify, resulting in a large carbon loss. This represents most of the carbon loss in a fluidized-bed gasifier.

In an overbed feed system, biomass particles are crushed to sizes less than 20 mm, which is usually coarser than the particle size used in the underbed system. In a typical setup, the fuel passes through bunkers, gravimetric feeders, and a belt conveyor, and is then dropped into a feed hopper.

Fewer feed points is an important characteristic of an overbed feed system. A rotary spreader throws the fuel particles over the bed surface. The coarser particles travel deeper into the gasifier while the finer particles drop closer to the feeder. The bed thus receives particles of a nonuniform size distribution. The maximum throwing distance of a typical spreader is around 4 to 5 m. The location of the spreaders is dependent on the dimensions of the bubbling bed. When the width is less than the depth, the spreaders are located on the side walls; when the depth is less than the width, they are located on the front wall. When both width and depth are greater than 4.5 m, the spreaders can be located on both side walls. Sometimes air is used to assist the throw of fuel by spreaders.

Tar is produced primarily through depolymerization during the pyrolysis stage of gasification. Biomass (or other feed), when fed into a gasifier, first undergoes pyrolysis that can begin at a relatively low temperature of 200 °C and complete at 500 °C. In this temperature range the cellulose, hemicellulose, and lignin components of biomass break down into primary tar, which is also known as wood oil or wood syrup. This contains oxygenates and primary organic con­densable molecules called primary tar (Milne et al, 1998, p. 13). Char is also produced at this stage. Above 500 °C the primary tar components start reform­ing into smaller, lighter noncondensable gases and a series of heavier molecules
4.2 Basics of Tar

called secondary tar. The noncondensable gases include CO2, CO, and H2O. At still higher temperatures, primary tar products are destroyed and tertiary products are produced.