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EPSRC THERMAL MANAGEMENT

OF INDUSTRIAL PROCESSES

Evaluation of the Biomass Drying Process

(January 2010)

Report Prepared by: SUWIC, Sheffield University

Researchers: Dr Hanning Li, Dr Qun Chen, Dr X Zhang, Dr K Finney

Investigators: Professor Jim Swithenbank Professor Vida N Sharifi

Sheffield University Waste Incineration Centre (SUWIC) Department of Chemical and Process Engineering Sheffield University

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Executive Summary

Very large amounts of low grade waste heat from the process industry is widely available in

cooling water at about 90°C and flue gases at 250°C to 400°C. Depending on the

temperature level of the heat available, these sources of waste heat could be used for a number

of applications such as district heating in cities, drying or torrification of biomass fuels, and

cooking in the food industry.

Dry or torrified biomass provides significant benefits for combustion and gasification,

mainly through fuel pre-treatment, increased boiler efficiency, lower flue gas emission and

improved boiler and gasifier operations. However, drying is a major heat consumption

process. An overall evaluation of the process is required before more detailed research,

design and construction work is carried out.

This report presents the results of an investigation in which the integration of the drying

process into a power station fuel system has been explored. In the process studied, waste

flue gas or hot water from a process industry plant with a waste heat output of 100MW is

utilized as the heating source for biomass drying. The dried biomass is provided as the input

fuel of a 40MW power generating station. The drying process is evaluated using either flue

gas or superheated steam as the heating source and a belt conveyor as the dryer. However,

this evaluation procedure can be readily adapted to other applications.

The evaluation is classified into three sections;

• The first one evaluates the thermal input requirement of the drying process, followed

by evaluation of the suitability of using the heat source from the waste flue gas or hot

water exiting from a 100MW industrial process plant. The drying process evaluated

in this section is based on the adiabatic drying process.

• In the second section, process equipment and the process are evaluated to determine

the size requirements of the drying equipment and associated equipment. According

to the size requirements, the costs are evaluated for the use of flue gas or superheated

steam as the drying media respectively. This report also evaluates the effect of

material choice on costs in the case using superheated steam as the drying medium.

• In the third section, various configurations of drying strategies using the flue gas

stream or the superheated steam stream are compared to evaluate their profitability.

Based on the results of these evaluations, it is concluded that waste flue gas or hot water

exiting from a 100MW industrial process plant could be used as the heating source for wood

drying for biomass combustion in a 40MW boiler. In the selection of flue gas or superheated

steam, flue gas usage would be lower capital cost but the environmental issue must be

considered. Superheated steam is a good option for fast drying, heat recovery and

environmental protection, but the high capital cost is an issue. In general, a 3-4 year

investment return is predicted by this investigation.

Thus in accordance with the grant proposal, Sheffield University has conducted an

extensive literature review on biomass drying and an evaluation of the drying process. This

report presents the results of a Case Study project in which integration of the drying process

into a power station fuel supply system has been explored based on our previous report that

contained a literature review of “Biomass Drying”. Various other sources of information

were used in order to compile this report. These included websites, journal publications,

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reports and communications with manufacturers and industry. The calculation methods are

based on data in published reports, books and papers. Microsoft Excel was used for

calculations.

Acknowledgements:

The authors would like to thank the Engineering and Physical Science Research Council

(EPSRC Thermal Management of Industrial Processes Consortium) for their financial and

technical support for this research work.

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List of Contents

Nomenclature 5

1. Introduction 8

2. Case Study: Description of Drying Systems 9

3. Mass Flow Rates Required by Drying 10

3.1 Capacity Estimation of Dryer 11

3.2 Estimation of Humidity in Flue Gas 12

4. Literature Review: Drying Mechanism 16

5. Capital Costs of Drying 22

6. Profitability 26

7. Conclusions 30

8. References 31

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Nomenclature

A, constant in Antoine equation [-]

Aeff effective area of drying [m2]

Adesign designed area of drying [m2]

Aheat exchanger heat transfer area of heat exchanger [m2]

B, constant in Antoine equation [-]

bh the price of the heat [€/W]

be the price of the electricity [€/W]

C, constant in Antoine equation [-]

or water concentration in chip [kg/m3]

CostD, drying cost [€]

CostDC, direct capital cost [€]

CostIDC, indirect capital cost [€]

CostRUN; running cost [€]

Costeq equipment cost [€]

CCapital, capital cost [€]

Ct cash benefit. [€]

Costm maintenance costs. [€]

Cp, air Specific heat of air [kJ/kg-K]

Cp, vapour Specific heat of water vapour [kJ/kg-K]

Cp, water Specific heat of water [kJ/kg-K]

Cp, wood Specific heat of wood [kJ/kg-K]

Deff effective diffusivity of water in wood [m2/s]

G, Lang factor [-]

Gf, mass flow rate of flue gas [kg/s]

Hw, Enthalpy of inlet and outlet biomass [kJ/kg]

Hf, Enthalpy of inlet and outlet flue gas [kJ/kg]

Hs, Enthalpy of inlet and outlet steam [kJ/kg]

Hwat Enthalpy of water [kJ/kg]

Hlatent latent heat of water [kJ/kg]

HC, humidity [kg-water/kg-air]

HCf, humidity of inlet and outlet flue gas [kg-water/kg-air]

H consum, the heat consumption [W]

HCfuel price of fuel [€/mwh]

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HCflue gas price of flue gas [€/mwh]

h, heat transfer coefficient [W/m2K]

L diffusion distance (half of the smallest chip dimension) [m]

Lextra extended belt length [m]

Mwood, Dry mass flow of biomass [kg/s]

Mflue, Dry mass flow of air [kg/s]

Msteam, Mass flow rate of steam [kg/s]

Mwat, Mass flow rate of water [kg/s]

MCw,in, Moisture of inlet and outlet biomass [kg-water/kg-wood]

NPV, Net present value [€]

P Pressure [atm]

Ps,in, Pressure of inlet and outlet steam [atm]

Psat Saturated Pressure [atm]

P consum the electricity consumption [W]

Q, thermal flow rate [W]

R Recycle ratio [-]

T Temperature [K] or [oC]

Ta Temperature of air [K]

Twood Temperature of wood [K]

Tw, Temperature of inlet and outlet biomass [K]

Tf, Temperature of inlet and outlet flue gas [K]

Ts, Temperature of inlet and outlet steam [K]

Twat, Temperature of water [K]

TDT total drying time [s]

t time [s] or [year]

U Bed heat transfer coefficient [W/m3K]

W evaporation rate of water [kg/s]

Wload unit loading of wood on the belt [kg/m2]

Wbelt belt width [m]

Wextra extended belt width [m]

x distance [m]

Z bed height [m]

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Greek symbols

ϕ relative humidity [-]

ρa density of air [kg/m3]

ρwood density of wood [kg/m3]

τ drying time [s]

or operating time [h/year]

τop total operating hours in one year [hour]

τwood residence time of wood in dryer [s]

Subscripts

in. inlet, initial

out, outlet, final

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1. Introduction

Biomass can be used as renewable energy for heat and power generation by combustion.

The term biomass refers to forest residues (such as dead trees, branches and tree stumps), yard

clippings, sawdust and wood chips, which can be utilised in energy production. Biomass is

usually burnt on a grate or in a fluidised-bed boiler to produce heat and electricity. The

moisture content of biomass typically varies between 50 and 63 wt% (water per total mass)

depending on the season, weather and type of biomass. The typical lower heating value (net

calorifc heating value) of dry bio fuels varies from 18.5 to 21kJ/kg (Kiranoudis, 1995).

Unfortunately, the energy needed for the evaporation of water in the combustion boiler cannot

be utilised in the power generation process, since the temperature level of the latent heat is too

low. However, a low level of fuel moisture could save a lot of the energy used in the

combustor for water evaporation,. It would also be beneficial in reduced boiler dimensions

and lowered unburned solids. Biomass with lower moisture could also reduce or eliminate

combustion control problems caused by fluctuations in moisture content.

This report investigates biomass drying for a 40MW power generation process. The

heating source assumed for the biomass drying consists of waste flue gas at 250oC – 450

oC

and 90oC hot water exiting from a 100MW industrial process plant.

Drying biomass requires the construction of drying plant. The process and equipment

design is a critical step for the construction. The core feature of the concept is to design a

suitable dryer and its associated equipment. Two alternative drying systems, flue gas drying,

and steam drying with a water pre-heating process, are investigated for cases where the key

boundary conditions of the drying process are different. A continuous belt dryer with a heat

exchanger (if steam drying is used) is considered as the dryer construction in both systems.

The dryer design mainly consists of the determination of various sizing and operational

variables. The evaluation of desired process variable values for each design is carried out by

using a suitable criterion, mainly based on economic considerations.

Process design for dryers has become an increasingly challenging problem, which aims at

the evaluation of; the proper type of equipment, its associated flow-sheet arrangement, its

optimal construction characteristics and the operating conditions involved in the overall

design. Most design work in this field encounters difficult problems related to the complex

drying conditions. Numerous theories have been developed for modelling the drying

processes aimed at optimum design. However, the thermo-physical properties and transport

coefficients in most models are only approximately known, thus producing inaccurate or

erroneous results on large scale industrial applications (Kiranoudis, 1995).

The main purpose of this report is to evaluate how the drying process should be carried out

in order to minimise drying costs. Both capital and running costs are included in the

evaluation. In this work, the conveyor-belt dryer design problem is investigated based on

product quality criteria combined with cost. The optimal construction and operational

process variables are evaluated by appropriately evaluating net profitable value (NPV). The

full set of efficient optimal solutions for the problem is evaluated and the effect of optimal

design variables on process construction and operational variables is investigated for a

characteristic industrial dryer design using wood chips.

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2. Case Study: Description of Drying System

In this investigation, the energy source with the greatest potential for biomass drying is flue

gas with a temperature of 350oC to 450

oC, which exits from a industrial process plant of

100MW. 60% of the plant output of low grade heat is thermally available as hot water at

90°C and the rest is hot flue gas. In this study, two process configurations are proposed and

compared. The first configuration is the direct usage of flue gas as a heating source. The

secondary one utilizes flue gas to raise the hot water to steam at a desired temperature. The

generated steam is then used as a drying energy source. The temperatures of the inlet gas

flows in the first drying option are in the range of 350-450oC as provided by the industrial

plant. The inlet superheated steam in the secondary option is typically 150–180oC at 1-2 bar.

The drying temperature is usually 20oC lower than the temperature of the heat source. The

use of flue gas and steam in the drying process is estimated based on an adiabatic drying

process. Fig. 2-1 illustrates the mass and energy balances of the adiabatic drying process

utilising flue gas as the direct heating source. Figure 2-2 presents the adiabatic drying

process heated by steam and its associated pre-heating process.

Figure 2-1 Mass and energy balance for adiabatic drying using flue gas as the direct heating

source.

Figure 2-2 Mass and energy balance for adiabatic drying using pressurised steam, and an

associated pre-heater.

Direct drying by flue gas can save the construction of a heat exchanger. Because the

drying temperatures are relatively low when a direct heat source of flue gas is used, the

dimensions of the dryer become larger than if higher drying temperatures are used. But,

higher drying temperatures will cause problems of volatile evaporation causing air emissions

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and higher fire risks. Steam used as drying medium represents, however, a more valuable

energy source than direct heat by flue gas, because by allowing the steam to expand in a

turbine it would be possible to get mechanical work (electricity) out of the steam process.

The higher capacity of absorbed water by the steam reduces the dimensions of the dryer, but

construction of the associated heat exchanger would increase the capital costs.

The flue gas or steam demand for drying depends on many drying parameters, and also the

type of dryer. The determination of flue gas or steam mass flow in continuous flow drying is

studied in this report. It is expected that the waste flue gas from the industrial plant could

provide enough heat energy for two alternative drying approaches.

The emissions released during drying are heavily dependent on the drying temperature [6].

Usually, the amount of emission increases considerably when the drying temperature is above

100oC. Because of the emissions, the exhaust air from the dryer could not be released into

the atmosphere, if the drying temperature is over 100oC. To eliminate the release of

emissions during drying, the exhaust air from the dryer can be conveniently added to

combustion air. If the drying temperature is below 100oC, it is usually possible to discharge

the exhaust air into the atmosphere. The outlet temperature of flue gas in the proposed

scheme is designed to be at temperatures below 100oC, based on a relative humidity of 90%

of saturated state.

3. Mass Flow Rates Required by Drying

Industrial conveyor-belt dryers are the most popular family of dryers for drying agricultural

products. They are equipped with a conveyor-belt on which the product to be dried is

uniformly distributed at the entrance. This study is focused on process investment for

conveyor-belt dryers for biomass drying. A typical flow sheet with a sketch of the interior of

the dryer is presented in Fig. 3-1.

Figure 3-1 Side view of a continuous cross-flow dryer

The determination of flue gas or steam mass flow in continuous flow drying is studied in

this section. As mentioned in the previous section, the energy source with the greatest

potential for biomass drying is flue gas at a temperature of 350oC to 450

oC, which is exiting

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from the industrial plant of 100MW. Table 3-1 and Table 3-2 list mass and volume flow

rates at various potential conditions.

Table 3-1 Flue gas exiting from a 100MW industrial process plant

Temperature (OC) 250 300 350 400 450

Mass flow rate (kg/s) 179.71 146.24 123.12 106.19 93.26

Vol. flow rate(0OC) (M3/s) 139.29 113.35 95.43 82.31 72.29

Vol. flow rate (M3/s) 266.84 237.91 217.77 202.90 191.44

Vol. flow rate (M3/h) 9.61E+05 8.56E+05 7.84E+05 7.30E+05 6.89E+05

Table 3-2 90OC hot water exiting from 100 MW industrial process plant

Mass flow rate (kg/s) 204.72

Mass flow rate (t/h) 737.00

The determination of mass flow rates in the two process configurations is based on

mathematical models of the dryer that involve heat and mass balances for the heating medium

and the biomass product streams, as well as heat and mass transfer phenomena that take place

during drying. The system of equations generated is subject to thermodynamic and

construction constraints that must be taken into consideration. The modelling and

calculations are based on the following steady-state operation assumptions:

• The dry mass flow of the solid fuel passing through the dryer is constant;

• The inlet velocity, temperature and moisture content of the heating medium are constant.

3.1. Dryer Capacity Estimation

The solid biomass fuel considered in this project will provide the energy input for a power

plant with an output power of 40MW. To supply the necessary energy from biomass, one

must assess the heating value and evaluate the mass flow rate of dry biomass. The heating

value (net calorific heating value) of dry bio fuels varies from 16 to 21 (kJ/kg) depending on

biomass type. White pine wood chips are the assumed biomass for this project. The

heating value of white pine is 16.664 (MJ/kg-d.b.). The mass flow rate of the dry biomass

can be calculated from the power input requirement and biomass heating value:

.)./(

)( (kg/s)M wood

bdkgMJvalueheatingfuel

MWplantpowerininputpower

−= (3-1)

The mass flow rate of the dry biomass supplied for combustion is also applicable for

selecting and designing a dryer. In the design and selection of a dryer, the evaporation rate

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of water from biomass is also an important parameter required to evaluate the capacity of the

dryer. The biomass throughput capacity of the dryer can be evaluated according to the mass

flow rate of solid biomass. The evaporation rate can be calculated from:

)MC-(MCMW(kg/s) outinwood ×= (3-2)

Here, MCin and MCout are respectively the biomass moisture content at the inlet and outlet

of the dryer. The evaporation rates of water from the biomass are also useful parameters for

evaluation of the designed dryer; such as mass flow rates of the drying medium, effective area

of the dryer, etc. Table 3-3 lists evaporation rates of water from the biomass with variation

of biomass moisture.

Table 3-3 Evaporation rates of water from the solid biomass

Moisture

change(wt%-wet)

initial, final (kg/s) (t/h)

0.6, 0.1 3.3339 12.0019

0.6, 0.2 3.0005 10.8017

0.6, 0.3 2.5718 9.2586

0.5, 0.1 2.1337 7.6812

0.5, 0.2 1.8003 6.4810

0.5, 0.3 1.3716 4.9379

3.2. Calculation of Flue Gas Humidity

It is common practice to assume that within the dryer interior, the temperature and humidity

of the air-drying stream follows an adiabatic process in order to estimate stream properties.

The adiabatic process is a thermodynamic procedure in which no heat is transferred to or

from the outside of the dryer. In the case of flue gas absorbing water from solid biomass,

one can assume that the enthalpy of the flue gas containing water is constant during the drying

process, This leads to the conclusion that the enthalpy of flue gas at the entrance to the dryer

is equal to that at the outlet of the dryer:

gas) fluedry -(kJ/kgHH outf,inf, = (3-3)

The enthalpy of the flue gas with water content can be approximately calculated by using

humidity data for air containing water:

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ffvapourpAirp HCHCCC ×+×+= latentf,,f HT) (H (3-4)

As the flue gas passes through the dryer, the flue gas will evaporate water in the solid

biomass into steam, leading to the increased humidity in flue gas. In the drying process

without extra heat supplied, i.e. adiabatic operation, the increased humidity in the flue gas is

coupled with decreased temperature in the flue gas. Meanwhile, the humidity in the flue gas

increases continuously until steam in the flue gas becomes saturated at an equilibrium

temperature. This is the maximum humidity in the flue gas. The saturation pressure and

temperature can be estimated by the Antoine equation:

TC

BA

+−=)(Plog sat10 (3-5)

T is temperature (oC). Psat is saturated pressure (mmHg). A= 8.07131; B=1730.63;

C=233.426. These values of parameters A, B, and C are employed below 100oC. The

saturated pressure represents the amount of steam existing in the flue gas. Hence, the

humidity in the flue gas can be calculated in terms of saturation pressure:

sat

sat

air

water

PP

P

MW

MW

ϕ

ϕ

−×=HC (3-6)

Here, ϕ is the relative moisture. Since the saturated state is difficult to achieve, the

maximum relative humidity can be estimated as 90%.

Equations (3-3) to (3-6) describe the calculations needed to evaluate the humidity and outlet

temperature in the flue gas. An iteration procedure is needed to accomplish the calculation.

As required by the design, a known amount of water should be removed from the solid

biomass, represented by water removing rate (kg-water/s). As a consequence, the

corresponding amount of flue gas must take the water vapour out of the dryer at the specified

humidity. The flow rate of flue gas is required to balance the mass of water evaporated from

solid material in the dryer. The mass flow rate of flue gas is determined according to the

water removing rates at a given humidity change in the flue gas:

)/(

)/(G f

airkgwaterkgHCHC

swaterkgrateremovingwater

outin −−−

−= (3-7)

The flow rate of flue gas can be also used for evaluating the availability of supplied flue gas

by the 100MW industrial process plant. Also, the flow rate of flue gas is used to check the

capacity of the dryer. Figure 3-2 shows the volumetric flow rates of flue gas with variation

of flue gas temperature at different moisture changes. The maximum flow rate of required

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flue gas is 2.3 x 105 m

3/h. The available flue gas supplies up to 9.6 x 10

5 m

3/h. The

supplied flue gas is therefore sufficient to dry the biomass if flue gas is the selected drying

medium. Figure 3-2 also demonstrates that a higher temperature of flue gas could reduce the

loading of flue gas in the dryer. For the same final moisture content of solid fuel, high initial

moisture content of solid fuel requires a higher flow rates of flue gas; and similarly, for the

same initial moisture content of solid fuel, a high final moisture of solid fuel requires a lower

flow rate of flue gas.

Figure 3-2 Flue gas flow rates vs flue gas temperature for drying wood

In the drying system using steam, superheated steam will be used to absorb water from solid.

In the absence of other gas species in the flue gas, water vapour is a major component in the

heating medium. The steam mass flow rate in the dryer is estimated based on an adiabatic

process limited by saturated steam at a given temperature. However, the parameters in

Antoine’s equation (3-5) are altered to apply at temperatures beyond 100oC. In the steam

drying process, the heating source is steam that will be partially generated by using the hot

water at 90oC. This integrated pre-heating process can be achieved by using the high

temperature flue gas heating as an additional source, as shown in figure 2-2. Thus, the flow

rates of flue gas are also required to be evaluated by a thermal balance for the energy required

to raise 90oC hot water to the desired superheated steam temperatures of 140-180

oC.

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Figure 3-3 Steam flow rates vs steam temperature for drying wood

Figure 3-3 shows the variation of mass flow rates of steam with steam temperature at

different moisture changes for steam generation using flue gas at 250oC. As expected, an

increased temperature reduces the flow rates of required steam. For the same final solid fuel

moisture content, a high initial moisture solid fuel requires a higher flow rate of steam; and,

for the same initial moisture of solid fuel, a higher final solid fuel moisture requires a lower

flow rate of steam. There is no effect of recycle ratio on required flow rates since the

recycled steam is mixed with generated steam before entering the dryer. However, variation

of recycle ratio significantly affects the operation of the pre-heater, because an increased

recycle ratio reduces the generation of steam leading to reducing the requirement for flue-gas

usage. Figures 3-4 and 3-5 demonstrate the effects of recycle ratios on the flow rates of flue

gas at different steam temperature and flue-gas temperature. It is easy to see that the

increased recycle ratio significantly reduces the use of flue gas. It is interesting to note that

at high recycle ratio, i.e. R=0.75, the flue gas usage is negligibly affected by either flue-gas

temperature for steam generation in the pre-heater or steam operation temperature in the

dryer.

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Figure 3-4 Flue gas flow rates required for generating steam at various steam

temperatures and steam recycle ratios

Figure 3-5 Flue gas flow rates required for generating steam at various flue gas

temperatures and steam recycle ratios

4. Literature Review: Drying Mechanism

In the operation of the drying process, air or steam as drying media flow through the solid

bed and contact with the surface of solid material. As a consequence, this convective drying

process removes water from the surface of solid material and increases the temperature of

solid because the temperature of stream is higher than of solid. During the convective

drying process, two periods can be distinguished in solid material (Gigler et al., 2000). In

the first so-called constant drying rate period, mainly external water attached to the product is

removed. In the second so-called falling drying rate period, internal diffusion of water to the

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surface of the product takes place. These physical phenomena have been described by two

types of models (Gigler et al. 2000; Fyhr and Rasmuson, 1997, Tang et al., 2004).

In the first kind of model, mass and energy balances are applied to describe convective

mass and heat transfer between the solid surface and stream. Of these models, partial

differential equations have a good physical basis to predict the drying process of a product

with appropriate drying models: given product mass, product characteristics and drying

stream conditions. In general, there are three equations needed to predict drying properties

of wood: stream temperature, wood temperature and drying rate based on mass and energy

balances.

For air drying, the simplified mass or water balance is based on the assumption that water,

evaporated by the wood [left-hand part of Eq. (4-1)], is taken up by the drying air (right-hand

part):

t

MC

x

Hwooda

∂

∂=

∂

∂ρρaM (4-1)

The simplified energy balance of the drying air describes the change of enthalpy of the air

[left-hand part of Eq. (4-2)] which is equal to the heat necessary to increase the temperature of

evaporated water from wood to air temperature, and the heat transfer from air to wood:

)()()(M ,,,a woodawoodavaporpwood

a

vaporpapa TTUt

MCTTc

x

THCcc −−

∂

∂−=

∂

∂+ ρρ (4-2)

The energy balance for the wood describes the change of enthalpy of the wood, which is

equal to the heat needed for evaporation of water and the heat transfer from the air to the

wood:

)()( ,, woodalatentwood

wood

waterpwoodpwood TTUt

MCH

x

TMCcc −−

∂

∂=

∂

∂+ ρρ (4-3)

Here, t

MC

∂

∂, also called the drying rate, represents the water transferred from the chip

surface to the drying stream. The drying rate is strongly dependent on the characteristics of

the wood, such as porosity, hardness, pore size, chip size, etc. Hence the drying rate of wood

is generally determined by experimental observations and then developed into an empirical

relation. Some researchers (Gigler et al. 2000; Fyhr and Rasmuson, 1997, Tang et al., 2004)

developed a theoretical approach to describe the drying rate based on the diffusion of water

from the interior of the chip to the surface, which can be expressed by a diffusion model for a

plane sheet

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)(x

CD

xt

Ceff

∂

∂

∂

∂=

∂

∂ (4-4)

One initial and two boundary conditions are necessary to fully describe the internal

diffusion process of a wood chip. The initial condition assumes that the original moisture

concentration of a wood chip is homogeneous. The first boundary condition describes

symmetry around the centre of the chip. The second boundary condition describes the mass

transfer of moisture to the air: thus the diffusive flux at the edge (x=±L) of a wood chip is

equal to the water removal from wood surface into the drying air.

As described above, the drying rate of wood is generally determined by experimental

observations that are then developed into an empirical relation. The experimental results of

drying rates, the variation of moisture with time, are the important data required to

accomplish the simulation. Meanwhile, the drying rate curve is also important for

determining the residence time of solid materials in the dryer. Some of the published

experimental results related to drying rates of wood are introduced as follows:

Gigler et al. (2000) carried out drying experiments for willows chips and simulated the

drying process for the same chips. Figure 4-1 shows their experimental results of variation

of moisture content with time for willow chips with different sizes. The system was

operated using air flow. At the beginning, convective heat transfer dominates the drying

process leading to a fast drop of moisture content with time. As the drying continues, the

water on the solid surface becomes less and internal diffusion of water inside the solid

becomes significant, leading to slowing down of the drying rate. Figure 4-1 also

demonstrates that an increased chip size retards the drying time. Hence the dryer requires a

longer drying duration to achieve the same final moisture of wood chips for drying larger

sizes of chips.

Figure 4-1 Drying curves for willow chips for a chip bed of 1 cm (●) and 8 cm (▲)

particles.

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Holmberg et al. (2007) studied the drying rates of pine wood bark with a dimension of

2x20x5 mm in a deep bed with variable heights using air as drying medium. Fig. 4-2 shows

the variation of moisture content with drying time for a variety of drying temperatures. As

shown in figure 4-2, an increased drying temperature significantly reduces the drying time

and hence accelerates the drying process for pine wood.

To evaluate the effect of bed height on air mass flow, the ratio between the drying residence

time and the bed height (τu/Z) was experimentally studied as shown in Fig. 4-3 (Holmberg et

al., 2007). Despite the variable final moisture content of the sample, the ratio τu/Z seems to

decrease constantly as the bed height increases. However, the derivative of the ratio is

clearly smaller when the bed heights are over 100-120 mm for inlet temperatures of 120oC.

Bed thicknesses approximately represent bed heights when the outlet air is fully saturated at

the beginning of drying. Holmberg et al.(2007) concluded that the bed must be at least so

high that the exit drying air reaches its saturation point at the beginning of the drying.

Increasing the bed height still decreases the dryer size but the influence on dimensions is not

as significant as for thin beds. Bed heights between 0.2 - 0.8m are generally selected for

conveyor dryers.

Figure 4-2 Drying curves obtained by Holmberg and Ahtila(2005)

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Figure 4-3 The ratio between drying time and bed height as a function of bed height

Fyhr and Rasmuson (1997) investigated the effects of different wood materials and

operational conditions in the superheated steam system. Comparing pine wood with spruce

wood, drying pine wood takes less time than spruce wood, as shown in Figure 4-4. This is

due mainly to the difference of internal structure that makes pine wood more permeable than

spruce. They also investigated the effects of solid sizes and operating temperature on drying

residence time, which can be approximately described by:

. o

o

oL

L

T

TTDTTDT 1

1

1 = (4-5)

Figure 4-4 Drying behaviour of spruce and pine woods

~ 21 ~

The variation of moisture with time in Figure 4-4 can be converted into a drying rate curve.

Figure 4-5 shows a typical drying curve. As shown in Figure 4-5, a maximal drying rate is

obtained at a shorter drying time. After that, the drying rate rapidly decreased. As

discussed above, the initial fast drying rate can be attributed to a convective drying process.

The slow drying rates at a later time can be attributed to internal diffusion of water in this

period.

Figure 4-5 depicts the effect of temperature on both superheated steam drying and air

drying. As the plot shows, the maximum drying rate is much higher with superheated steam

than with relatively dry air at temperatures above approximately 180oC, while the relationship

is reversed below this point. The maximum drying rate represents initial drying rate, which

will identify the optimal operation conditions. In view of the drying rate, air drying will be a

preferred option to accelerate the drying process at drying temperatures below 180oC, whilst

steam drying will significantly improve the drying rate at temperatures beyond 180oC.

Figure 4-5 Typical drying rate curve

~ 22 ~

Figure 4-6 Maximum drying rate vs temperature in air and steam

5. Capital Costs of Drying

In the design of a new process, cost evaluation plays an important role by helping

investigators to make the right decision. The classification of costs is mainly based on the

principles presented in Brennan (1998). The costs generally consist of both capital and

running costs. Capital costs are usually divided into direct and indirect costs. The total

drying costs CostD may be written as follows

CostD = CostDC + CostIDC + CostRUN; (5-1)

where CostDC represents direct capital costs, CostIDC indirect capital costs, and CostRUN

running costs.

Direct capital costs are calculated by multiplying purchased equipment costs by a given

factor, defined as the “Lang factor”(Brennan, 1998)). Purchased equipment costs are

frequently presented in chart form, and they are usually correlated with a capacity factor using

the relationship:

bkY=eqCost (5-2)

~ 23 ~

where k is the proportionality factor, Y the capacity parameter and b the exponent. Exponent

b is typically within a range of 0.4–0.8 (Brennan, 1998) and leads to the economy of scale.

In drying systems, the main pieces of equipment are conveyors, heat exchangers, and fans.

The capacity factor of each piece of equipment is different from each other, but it is directly

or indirectly dependent on the dry mass flow of drying air or steam.

The Lang factor is the sum of several factors which are applied for the estimation of costs,

such as instrumentation, electrical, erection, structures, and lagging. The value of each

individual factor depends on the purchased equipment costs. Some approximate values for

these factors are listed in (Brennan, 1998). The direct capital costs of the dryer may be

written as follows:

CostDC = G Costeq (5-3)

G represents the Lang factor in (5-3).

Indirect costs cover engineering and project management, as well as a contingency

allowance, which can be considerable in pilot plans. Indirect costs are usually added as a

percentage of direct capital costs, and they are not dependent on the dimensions of the dryer.

In drying, running costs encompass all those costs associated with the operation of the dryer.

The most important running costs are composed of the use of heat and electricity and

maintenance costs. The costs covering the use of heat and electricity are dependent on the

annual operation time of the dryer and the price of energy. Maintenance costs are usually

estimated as a percentage of direct capital costs, and typical values range from 2 % to 11%,

averaging around 5% to 6% (Brennan, 1998). Personnel costs and insurance are also

included in the running costs. They are, however, heavily site dependent and therefore more

difficult to define. If the operating time of the dryer is τ h/year the annual running costs

become

CostRUN = Hconsum τ bh + Pconsum τ be + Costm + Costx; (5-4)

where Hconsum is the heat consumption (W), Pconsum the electricity consumption (W), bh the

price of the heat, be the price of the electricity, and Costm maintenance costs. The price of

the heat depends on the heat source. The main consumer of electricity is fans. The heating

used in this project is either flue gas or by-product hot water from an industrial process plant,

hence, the prices of heat consumption (and electricity) may not need to be accounted for.

The term CostX represents all other running costs (e.g. personnel, and insurance).

Capital cost is generally considered as a direct cost, because of the construction and

equipment costs, and the other costs are related to the capital costs. As mentioned above, the

capital costs consist of equipment and construction costs:

∑= ieqCostG ,DCCost (5-5)

~ 24 ~

Here, G is selected as 1.6, including 0.1 for electricity, 0.1 for instrumentation, 0.05 for

lagging, 0.15 for civil work and 0.2 for installation.

The cost functions of individual equipment items are summarized as follows:

Belt dryer: Costeq=2700Y Y is cross-section area (5-6)

Heat exchanger: Costeq=660Y0.7

Y is heat transfer area (5-7)

Cover: Costeq=1200Y0.5

Y is cover area (5-8)

All of these costs are counted in either € or converted into €. In the cost of the belt dryer,

the important capacity parameter is belt cross-section area. To obtain the area, solid mass

flow rate and residence time for drying should be known. Solid mass flow rate (Mwood) can

be evaluated based on the data in the previous section. The residence time (τwood ) is

considered as pine wood drying time as shown in Figure 4-2. As the total wet solid amount

on the belt is known (Mwood *(1+MC)*τwood), the required belt area can be estimated based on

unit area loading of wet solid. Pareso(2002) recommended a maximum unit area loading of

50kg/m2. A unit area loading of Wload = 30kg/m

2 is used in this report. The effective area is

estimated as:

load

woodwood

W

MCM τ×+=

)1(A eff (5-9)

An extra length and width of belt may be added for end effects. The designed area can

then be estimated as:

)()(A design extrabeltextra

belt

effWWL

W

A+×+= (5-10)

The equipment cost of the conveyor can be calculated by using Adesign as the cross-section

area in cost function (5-6)

The equipment cost of the cover is based on the designed area that covers the conveyor-belt.

The designed length and width of cover may be slightly longer and wider than of the belt.

The height of cover is designed as 6 meters as that height is usually used for an industrial belt

dryer.

In the calculation of equipment cost for the heat exchanger, the heat exchange area is

evaluated based on convective heat transfer equations:

)(A exchangerheat

watf TTh

Q

−×= (5-12)

The heat exchanger in this study is used to convert hot water at 90oC into steam at desired

~ 25 ~

temperatures of 140-180oC. During this process, the heat transfer parameter Q consists of

three sections, water temperature rise from 90oC to 100

oC, water evaporation to steam at

100oC, and steam temperature increasing to the desired temperature. Tf is the average value

of inlet and outlet flue gas temperatures for the heat exchanger. Twater is the water

temperature. Because water evaporation to steam at 100oC is heavily energy intensive, Twater

is set to 100oC

Figure 5-1 shows the variation of capital costs with solid final moisture at an initial

moisture content of 1.5 (kg-water/kg-solid dry) under different flue gas temperatures. As

expected, process construction for drying to a higher final solid moisture level and higher

operation temperature could reduce the capital cost.

Figure 5-1 Variation of capital cost with solid final moisture at an initial solid moisture 1.5

(kg-water/kg-solid dry) for different flue gas temperatures

Because the operation with steam drying is likely causes corrosion problems, stainless steel

material is partially used for equipment construction. For the heat exchanger, tubes are

constructed of stainless steel and the shell is constructed of carbon steel. The dryer is

manufactured using either carbon steel or stainless steel. Figure 5-2 shows the variation of

capital costs with solid final moisture at an initial moisture content of 1.5 (kg-water/kg-solid

dry) for different steam conditions and equipment materials. As expected, a drying process

for higher final solid moisture and higher operation temperature could reduce the capital cost.

High pressure steam operation generally causes a corrosion problem with the equipment,

however stainless steel material would reduce the problem. Nevertheless, the capital cost

will be significantly increased, as shown in Figure 5-2.

~ 26 ~

Figure 5-2 Variation of capital cost with solid final moisture at an initial solid moisture

content of 1.5 (kg-water/kg-solid dry) under different steam conditions and equipment

materials

6. Profitability

Based on the cumulative cash flow, the profitability can be evaluated in terms of time, cash

and percentage return on investment. Payback time is generally the main concern for

investors. It is sometimes taken as the time from commencement of the project to recovery

of the initial capital investment. More frequently, it is taken as the time from the start of

production to recover the fixed capital expenditure only; working capital and land capital are

considered as being recoverable at the conclusion of the project.

When measuring profitability, the change of money over time must be accounted for. Net

present value (NPV) is a measure of the net cash benefit generated by the project. This

report utilizes NPV to evaluate the profitability of the designed drying processes.

Capital

enancema CC

−+

−= ∑

=

=

kt

0tt

intt

i)(1

C t)NPV(projec (6-1)

Here, CCapital and Cmaintenance are capital and maintenance costs as discussed above. i is the

interest rate. t is an individual year. k is total number of years. Ct is cash benefit in t

years.

Cmaintenance=0.05 CCapital (6-2)

~ 27 ~

Ct= (HCsave-Q•HCflue gas) • τop (6-3)

τop (8400h) is total operating hours in year t. Price of flue gas, HCflue gas, is 0-0.5 €/GWh.

Here 0.5 €/GWh was used for the price of flue gas. If heating rate for water evaporation is Q

(kJ/s), flue gas cost would be Q• HCflue gas.

HCsave is defined as hourly price in saved fuel. When the water content in biomass is

reduced, i.e. reduced from 1.5 to 0.1 water-weight/dry solid, the heating value (HHV) of the

biomass will be increased. The increased HHV will be beneficial in saving energy on boiler

operation to evaporate the same amount of water as that removed from the solid in drying.

The saved energy in the boiler can be converted into a positive cash flow as follows:

fuellatent HCH ××= waterevaporatedSave WHC (6-4)

HCfuel is the price of fuel. An item of heatlatentH× waterevaporatedW represents the total energy

required to evaporate the desired amount of water in one hour. HCfuel, fuel price, is generally

dependent on the type of fuel, time and other parameters; but, it could be reasonably

considered as the same price as biomass fuel used in the drying-boiler process. Currently,

the biomass cost has not been specifically selected. The fuel price, HCfuel, is generally priced

in the range of 6-20 (€/MWH). In the calculations, 14 €/MWH is used for the fuel price

used in the following discussion.

Figure 6-1 shows the variation of NPV with operation at an operating temperature of 90oC.

The drying process is designed for drying white pine from an initial moisture content of 1.5

kg-water/kg-solid to finial moisture levels of 0.1 and 0.3 (kg-water/kg-solid). If flue gas is

used for drying the biomass, the return on investment should be achieved after 3-4 years of

operation. Figure 6-2 compares the 10-year NPV at different solid moisture levels of the

product. At an operation temperature of 110oC, a profit of 3.6 million € can be reached with

product moisture as low as 0.1 kg-water/kg-solid. An increased product moisture content

lowers the profit. At an operation temperature of 90oC, the most profitable value is found at

product moisture of 0.3 kg-water/kg-solid.

~ 28 ~

Figur 6-1 Cumulative cash flow for biomass drying by flue gas with an operating

temperature of 90oC and drying biomass from an initial moisture of 1.5 kg-water/kg-solid to

0.1 and 0.3 kg-water/kg-solid.

Figur 6-2 Variation of NPV at 10 year with final solid moisture

(the initial moisture: 1.5 kg-water/kg-solid; flue gas).

Figure 6-3 shows the cumulative cash flow of biomass drying using superheated steam at an

~ 29 ~

operating temperature of 150 oC. As shown in the figure, 3-4 years of operation is expected

to be needed to achieve the return on investment. Figure 6-4 shows the cumulative cash

flow up to 10 years at an operating temperature of 150oC. In general, NPV is slowly reduced

with final moisture until the moisture is up to 0.25 and then the profitability quickly drops

with increased final moisture. This indicates that the dried biomass would yield a higher

economic benefit. Figures 6-1 to 6-4 show the evaluated results based on a fuel price of 14

€/MWh. As expected, fuel price is a sensitive factor to achieve profitability. Figure 6-5

shows the effect of fuel prices on NPV at 10 years. As shown in the figure, a higher biomass

fuel price brings about higher profitability. Figure 6-5 also demonstrates that the required

fuel price should be higher than 8 €/MWh for a positive return on investment at 10 years.

Figur 6-3 Cumulative cash flow for biomass drying by superheated steam at an operating

temperature of 150oC and drying biomass from an initial moisture content of 1.5

kg-water/kg-solid to 0.1 kg-water/kg-solid.

~ 30 ~

Figur 6-4 Variation of NPV at 10 years with final solid moisture where the initial moisture

is 1.5 kg-water/kg-solid; 150°C steam.

Figur 6-5 Variation of NPV at 10 year with fuel price (the initial and

final moisture: 1.5 and 0.1 kg-water/kg-solid; 150oC steam).

7. Conclusions

In this report, integration of the drying process into an industrial process plant with a power

station system has been studied. In the process investigated, waste flue gas or hot water

~ 31 ~

exiting from a 100MW industrial plant is utilized as the heating source for biomass drying;

the dried biomass is provided for the input of a 40 MW power generation station. The

drying process is evaluated using either flue gas or superheated steam as the heating source

with a belt conveyor as dryer.

Based on the result of these evaluations, waste flue gas and hot water exiting from a

100MW industrial process plant could be used as the heating source for wood drying for

biomass combustion in a 40MW power station boiler.

By using flue gas drying, capital cost would be 2.5 million €. A higher flue gas

temperature would reduce the capital cost; but environmental issues will be a concern. By

using steam drying, capital cost would reach 3 million €. To protect corrosion of equipment,

stainless steel material can be introduced into the construction of equipment, but the cost will

be doubled. So, in the selection of flue gas or superheated steam, flue gas usage could give

lower capital cost but the environmental issue should be considered. Superheated steam

should be a good option for fast drying, heat recovery and environmental protection, but high

capital cost is an issue, particularly in the partial usage of stainless material for the equipment.

In general, for both flue gas and steam drying, 3-4 year operation is expected to give a

return on the investment at a fuel price of 14 €/MWH. However, biomass fuel price is a

sensitive factor to achieve profitability for biomass drying. The price should be higher than

8 €/MWH to achieve a return on the investment after 10 years of operation.

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