Removal of NH3 from Biomass Gasification Producer Gas by Water

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Ind. Eng. Chem. Res. 2005, 44, 1576-1584

Removal of NH3 from Biomass Gasification Producer Gas by Water Condensing in an Organic Solvent Scrubber Tobias Pro1 ll,* Ingmar G. Siefert, Anton Friedl, and Hermann Hofbauer Institute of Chemical Engineering, Vienna University of Technology, Getreidemarkt 9, 1060 Wien, Austria

Producer gas from biomass gasification contains NH3, H2S, and high molecular weight organic compounds (tars), which must be removed prior to the gas utilization step. A highly effective approach toward the removal of tars is absorption into organic solvents. As a secondary effect, condensation of water takes place in the scrubber, allowing removal of gaseous trace components such as NH3 and H2S from the gas stream. The present work focuses on the coabsorption of NH3 and CO2 into condensing water. The general aim is to decrease the clean gas NH3 concentration by appropriate adjustment of the operating parameters without additional gascleaning steps. A mathematical model of a countercurrent column, which covers mass and energy balances, gas-liquid equilibrium, and electrolyte dissociation, is presented. The scrubber geometry does not enter the model; therefore, it promises universal applicability. Property data on electrolyte dissociation and Henry’s law constants are taken from the literature. A computer code has been developed in order to study the sensitivity of NH3 removal toward certain process parameters. The results agree well with data measured at the 2000 Nm3/h (gas load) scrubber at the biomass gasification plant in Guessing, Austria. 1. Introduction

Table 1. Typical Composition of a Producer Gas from the FICFB Gasification Process before Gas Cleaninga

Biomass steam gasification allows the conversion of solid biomass to medium calorific gas consisting mainly of hydrogen, carbon oxide, methane, carbon dioxide, and water. Either the gas may be used for electricity production in turbines or engines or it may be used for the synthesis of high-quality liquid fuels or synthetic natural gas. In the fast internal circulating fluidized bed (FICFB) gasification process,1,2 the solid fuel enters a steamfluidized stationary bed where thermal conversion takes place at temperatures between 850 and 900 °C. The energy for the endothermic gasification reactions is introduced by the circulating-bed material heated in a second fluidized-bed reactor by combustion of residual char. Gas cleaning is necessary in order to remove particles and high molecular weight organic compounds (tars) from the producer gas. An 8 MW (fuel power) combined heat power (CHP) plant, coupling an FICFB biomass steam gasifier and a 2 MWel gas engine, has been realized in Guessing, Austria, and has been in operation since 2002. The plant uses wood chips as fuel. The gas-cleaning technology consists of a bag filter for the removal of particulates and an organic solvent scrubber for tar removal. The process is described in detail by Hofbauer et al.3 Table 1 shows the typical composition of the producer gas of the Guessing plant if untreated wood chips are used as fuel. The water content is strongly influenced by the ratio between total water (fluidization steam + fuel humidity) and dry fuel in the gasifier (steam/fuel ratio). Apart from solids and tars, there are also gaseous trace components in the producer gas, which may cause problems in gas utilization. One of them is NH3, which occurs in concentrations up to 1700 vppm (dry) in the raw producer gas. In the case of gas utilization in the combustion engine, NH3 in these concentrations reduces the operating lifetime of the engine oil and leads to

H2O CH4 C2H4 C3 fraction CO CO2 H2 N2 H2S NH3 tar particulates LHV

30-45 vol % 10-11 vol % (dry) 2-2.5 vol % (dry) 0.5-0.7 vol % (dry) 24-26 vol % (dry) 20-22 vol % (dry) 38-40 vol % (dry) 1.2-2.0 vol % (dry) 130-170 vppm (dry) 1100-1700 vppm (dry) 2-5 g/Nm3 (dry) 20-30 g/Nm3 (dry) 12.9-13.6 MJ/Nm3 (dry)

a All quantities except the water concentration are related to the dry gas fraction.

higher NOx emissions in the exhaust gas. For alternative gas applications such as Fischer-Tropsch synthesis for liquid fuel production, the required NH3 limits are as low as 1 vppm.4 The nitrogen content of wood for the different species is 0.10-0.45 wt % (waf).5 Nitrogen is concentrated in the bark. Soft wood generally contains less nitrogen than hard wood. Waste wood may contain much more nitrogen, up to 3.5 wt % (waf).5 According to the nitrogen balance for the Guessing plant, up to 70% of the fuel nitrogen is found as NH3 in the producer gas, even though the equilibrium of the inverse ammonia synthesis reaction

2NH3 S N2 + 3H2

(1)

is far on the side of nitrogen and hydrogen under gasification conditions [∼4 vppm (dry) of NH3 in equilibrium at 880 °C; 1 bar]. Apparently, kinetics does not allow significant conversion of NH3 to N2. After leaving the gasifier, the producer gas is cooled in an indirect heat exchanger to a bag filter inlet temperature of about 170 °C, freezing the NH3 concentration at the value of

10.1021/ie049669v CCC: $30.25 © 2005 American Chemical Society Published on Web 01/26/2005

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heat-exchanger entry. Up to now, no effects of the bag filter on NH3 in the producer gas could be observed. The organic solvent scrubber downstream of the filter cools the producer gas again to about 40-50 °C (with the temperature changing with ambient conditions), which implies that the main part of the water is condensing here, forming an emulsion with the organic phase. The organic solvent used is rape oil methyl ester (RME), which shows excellent tar solubility of up to 0.5 kg/kg at 50 °C. Experimental investigations have shown that NH3 is not significantly soluble in RME. For the present work, it is assumed that only the condensate accounts for NH3 removal in the tar scrubber. The scope of the present work is to gain a deeper insight in the mechanism of NH3 removal in the organic scrubber. Implying that the organic solvent is not interacting, the gas-liquid equilibrium between the condensate and producer gas in the scrubber is of prior importance for NH3 removal. The equilibrium in the liquid phase is governed by the interaction of the two weak electrolytes, NH3 and CO2. The behavior of aqueous solutions of weak electrolytes has been extensively investigated by Edwards et al.,6 who present a model for the system NH3-CO2-H2O based on the reactions

H2O S H+ + OH-

(2)

NH4+ S H+ + NH3

(3)

CO2 + H2O S H+ + HCO3-

(4)

HCO3- S H+ + CO32-

(5)

HCO3- + NH3 S NH2COO- + H2O

(6)

Figure 1. Dissociation constants of the NH3-CO2-H2O system: constants defined according to eqs 19-22.

The equilibrium at the gas-liquid interface is described by Henry’s law for the two species NH3 and CO2:

pA ) KH,Aaa

(7)

Equilibrium data covering the whole NH3-CO2-H2O system are presented by Edwards et al.7 These authors apply the correlations for water and ammonia dissociation reported by Tsonopoulos et al.8 (Figure 1). The water dissociation constant by Tsonopoulos et al.8 is in very good agreement with other literature sources9 for the temperature range 0-225 °C. The correlation for ammonia dissociation by Tsonopoulos et al.8 agrees well with the correlation by Kawazuishi and Prausnitz10 up to 150 °C. Henry’s law constant for NH3 by Edwards et al.7 (Figure 2) agrees with the correlation by Clegg and Brimblecombe11 at low temperatures and fits DECHEMA data12 at intermediate temperatures up to about 150 °C. The correlation of Kawazuishi and Prausnitz,10 who refer to data measured by Gillespie et al.13 and Mu¨ller et al.,14 yields a slightly lower solubility of NH3 in water than the Edwards correlation (difference of 7% at 50 °C) and is recommended up to 300 °C. Comprehensive data on Henry’s law and dissociation constants for the CO2-H2O system up to 250 °C are reported by Plummer and Busenberg,15 who suggest a five-parameter approach for the temperature dependence. The correlation for Henry’s law constant of CO2 determined by Chen et al.16 agrees well with the Plummer and Busenberg formulation. The equilibrium constant of carbamate formation (eq 6) is reported by Edwards et al.,6,7 Pawlikowski et al.,17 and Kawazuishi

Figure 2. Henry’s law constants of NH3 and CO2 in water and equilibrium constant of carbamate formation: constants defined according to eqs 17, 18, and 23.

and Prausnitz.10 The correlations differ by up to 50% between 30 and 100 °C, where the correlation of Pawlikowski et al.17 differs also in shape and shows a peak at about 50 °C. Within the present work, the newest correlation by Kawazuishi and Prausnitz,10 recommended up to 160 °C, has been chosen. Anyway, reaction (6) influences the total NH3 solubility only marginally within the temperature and pH range of the present work.6 2. Experimental Section Setup of the Organic Solvent Scrubber. Figure 3 shows the setup of the tar scrubber with a capacity of

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with 350 mL of 0.05 M H2SO4 each. The dry gas flow is monitored with a gas meter. One sampling run lasts for 30 min with a sampling flow rate of 0.080 Nm3/h. The solvent of the four impinger bottles is put together for further analysis. Condensate Sampling. During the gas sampling run, three samples of condensate are taken from the exit of the equalization tank (Figure 3). The three samples are merged to get the condensate representative of the gas compositions. Determination of NH3 in Liquid Samples. The total amount of ammonia in the liquid samples is determined by ion chromatography in Dionex DX-120. Because the samples are introduced by an acid buffer solution, NH3 is quantitatively detected as NH4+. The basicity of the condensate has been determined at room temperature with a pH meter (Mettler-Toledo pH 2-14 electrode). 3. Mathematical Model Figure 3. Scrubber system at the biomass gasification CHP plant in Guessing, Austria.

Figure 4. Gas sampling setup.

2000 Nm3/h (dry) in the biomass gasification CHP plant in Guessing, Austria. The ratio of the new organic solvent to the circulating solvent is about 1:2500 in practical operation. The temperatures vary according to the ambient conditions (winter/summer). Temperatures and flow rates are automatically registered during operation. The producer gas CO2 content is measured twice: once by an infrared online analyzer (Advanced Optima by ABB Automation) and, together with other gas components, also by an online gas chromatograph (Syntech Spectras GC 955; 80 °C operating temperature; He as the carrier gas; two columns, Chromosorb for splitting between permanent gases, CO2, and hydrocarbons and a 5 Å molecular sieve for splitting the permanent gases H2, O2, N2, CO, and CH4). The raw gas water content can be calculated from a water balance across the scrubber because the amount of condensate can be directly measured, and the clean gas water content can be calculated from the gas exit temperature assuming 100% saturation. Ammonia Sampling from Gas. The measurement principle is based on the discontinuous sampling of the producer gas. The gas sampling points are situated before and after the scrubber (refer to Figure 3). The sampling system consists of a heated probe with a spherical valve and a series of cooled impinger bottles containing an acid solution to catch the ammonia. The temperature of the cooling agent is -2 °C, so that the gas is cooled to almost 0 °C and practically free of water (Figure 4). To get representative values before and after the scrubber, raw and clean producer gases are extracted simultaneously with separate sampling systems according to Figure 4. The cooled impinger bottles are filled

The model aims at the description of the phenomena relevant for NH3 removal inside the scrubber. As mentioned in the Introduction, NH3 has not been found to be soluble in the organic phase and can therefore only be absorbed by the condensing water. Because of the complex behavior of the two liquid phases, no reliable correlation for the mass-transfer coefficient between the gas and aqueous phases could be found. On the other hand, phase separation experiments show that the organic phase tends to cover the gas-liquid interface, which leads to the major assumption in the model: It is assumed that water is in equilibrium with the gas phase just when condensation takes place and then immediately emulsified in the organic phase. Once emulsified, any mass transfer from condensate to gas is inhibited by the organic phase. The condensate in the equalization tank represents a mixture of the water fractions occurring along the height of the column. Energetic Aspects of Condensation. If the clean gas exit temperature is below the saturation temperature of the raw gas mixture, the partial pressure of water in the clean gas will equal the vapor pressure of water at the exit temperature. The rest of the water introduced with the raw gas leaves the scrubber as liquid emulsified in the organic phase. All aspects of tar removal are neglected in the modeling. The column is modeled one-dimensionally and without axial mixing in any of the phases. This simplified description of phase flow together with the assumption of inhibited mass transfer introduced above allows the utilization of the gas-phase temperature as a coordinate instead of the geometrical height. Therefore, the model does not depend on the column geometry. In the lower part of the column, only heat transfer from gas to emulsion occurs until the gas phase reaches the saturation temperature and condensation starts. No reevaporation of emulsified condensate is considered in the lower part of the column. The gas phase is modeled as an ideal gas, which means that the product of the total pressure and gas-phase mole fraction yi is proportional to fugacity and that the total gas enthalpy is a linear combination of the enthalpies of the various components. The saturation gas temperature TG,C can be calculated by eq 8 with the raw gas water content yH2O,in and total pressure p. Once condensation has started, the water content in the gas decreases, according to eq 8, with decreasing


pyH2O,in ) pS,H2O(TG,C)

(8)

TG. The saturation vapor pressure of water pS,H2O is given in bar by the DIPPR 801 equation,18 which fits the IAPWS-IF9719 data in the whole temperature range in which liquid water may appear (273.16-647.13 K).

ln(pS,H2O) ) 62.1361 -

7258.2 - 7.3037 ln(T) + T 4.1653 × 10-6T 2 (9)

For the formulation of mass and energy balances, it is advantageous to relate all extensive quantities to the flow rate of dry gas. The water load on dry gas is

Y H 2O )

yH 2 O

(10)

1 - yH2O

The upper part of the column, where condensation occurs, can be divided into n temperature intervals ∆T of equal size. The specific amount of condensate ∆YCOND,k occurring in element k can be gained from eq 11, with the indices “in” and “out” referring here to element k.

∆YCOND,k ) YH2O,k,in - YH2O,k,out

(11)

The gas-phase water loads YH2O,k,in and YH2O,k,out can be calculated from eqs 8-10 with the temperatures at the boundaries of element k. The total specific amount of condensate YCOND is determined by the global water balance. n

YCOND ) YH2O,in - YH2O,out )

∑ ∆YCOND,k

(12)

k)1

The indices “in” and “out” refer here to the gas streams entering and exiting the scrubber. Neglecting heat losses, the global energy balance can be written as I (TL,out - TL,in) + YCONDHWAT(TL,out) + CSOLV HGAS,dry(TG,out) - HGAS,dry(TG,in) + YH2O,outHGAS,H2O(TG,out) - YH2O,inHGAS,H2O(TG,in) ) 0

(13) TL refers to the liquid-phase temperature, which is uniform for organic and aqueous phases. YCOND is the specific amount of condensate per mole of dry gas, YH2O is the H2O load on dry gas, and H(T) refers to the conventional enthalpy (sensible heat including the enthalpy of formation). The indices “in” and “out” refer to whether a stream enters or leaves the system, “WAT” indicates liquid water, “GAS,H2O” indicates gaseous water, and “GAS,dry” indicates the dry gas composition. The mean heat capacity of the organic solvent CISOLV is defined with the specific heat capacity cP,SOLV, the mass flow of solvent m ˘ SOLV, and the molar flow of dry gas n˘ G,dry as follows: I CSOLV

)m ˘ SOLVcP,SOLV/n˘ G,dry

(14)

If the initial gas composition and the in/out temperatures of the gas and liquid phases are known, CISOLV can be calculated from eq 13. When the mean heat

capacity of the organic solvent is treated as a constant within the column, a partial heat balance in the form of eq 13, starting either from the top or from the bottom of the column, allows the calculation of the corresponding liquid-phase temperature to any gas temperature between TG,in and TG,out. An interesting parameter to check the plausibility of model assumptions and measured data during model evaluation is the minimum temperature difference ∆TGL,C between the gas and liquid phases in the section where condensation starts.

∆TGL,C ) TG,C - TL,C

(15)

The partial heat balance in the lower part of the scrubber for the calculation of TL,C can be written as I (TL,out - TL,C) + YCOND[HWAT(TL,out) CSOLV HWAT(TL,C)] + HGAS,dry(TG,C) - HGAS,dry(TG,in) + YH2O,in[HGAS,H2O(TG,C) - HGAS,H2O(TG,in)] ) 0 (16)

“C” indicates the section where condensation starts. Because the total pressure is approximately 1 bar, the producer gas is assumed to behave like an ideal gas and ideal gas enthalpy data are used. The pressure dependence of the enthalpy of liquid water can also be neglected. Burcat and McBride20 report coefficients for the calculation of total enthalpies H0(T) for various species using the NASA polynomials. The molar enthalpy of ideal gas mixtures is the linear combination of the enthalpies of the different components using the mole fractions as weights. NH3 Absorption. From the section where condensation starts upward, water is condensing in chemical equilibrium with the gas phase and transported downward as an emulsion in the organic phase. Because mass transfer between the gas and aqueous liquid phases is assumed to be inhibited by the organic phase, the lower part of the column, where no condensation occurs, does not contribute to NH3 absorption. Gas-liquid equilibrium during condensation is calculated according to the reaction system presented by Edwards6 (eqs 2-6). Because of the low electrolyte concentrations in the condensate, the solution is assumed to behave ideally. In other words, all activity coefficients in the solution are assumed to be unity and the equilibrium is expressed by molalities m (moles per kilogram of water). The set of equations to describe the gas-liquid equilibrium is presented as follows.

KH,NH3 ) pyNH3/mNH3(aq)

(17)

KH,CO2 ) pyCO2/mCO2(aq)

(18)

Kw ) mH+mOH-

(19)

KNH4+ ) mH+mNH3(aq)/mNH4+

(20)

KCO2 ) mH+mHCO3-/mCO2(aq)

(21)

KHCO3- ) mH+mCO32-/mHCO3-

(22)

KNH2COO- ) mNH2COO-/mNH3(aq)mHCO3-

(23)

mH+ + mNH4+ ) mOH- + mHCO3- + 2mCO32- + mNH2COO- (24)

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Table 2. Effect of Temperature on Equilibrium Constants According to Equation 25 KH,NH3 [bar kg/mol] KH,CO2 [bar kg/mol] Kw [mol2/kg2] KNH4+ [mol/kg] KCO2 [mol/kg] KHCO3- [mol/kg] KNH2COO- [kg/mol]

A0

A1

A2

A3

41.79296 -108.3808 61.2062 60.0072 -356.3094 -107.8871 8.75196

0.003 733 62 -0.019 850 78

-13.588 57 40.451 54 -22.477 3 -23.974 4 126.833 9 38.925 61 -4.017 263

-3291.9296 6919.53 -5839.5 -4390.82 21834.37 5157.79 262.3644

0.016 093 5 -0.060 919 64 -0.032 528 49 0.002 184 91

Table 2 reports literature coefficients for the temperature dependence of the equilibrium constants according to the equation

log(K) ) A0 + A1T + A2 log(T) +

A3 A4 + 2 (25) T T

with T in Kelvin. For known temperature and gas molar fractions of NH3 and CO2, the solution is determined by eqs 1724. The total amount of electrolytes in the liquid is then

mNH3(total) ) mNH3(aq) + mNH4- + mNH2COO- (26) mCO2(total) ) mCO2(aq) + mHCO3- + mCO32- + mNH2COO(27) The total amount of NH3 or CO2 in the condensate leaving the column can be calculated with Tk as the medium temperature in element k as follows.

mi(total),out )

1 YCOND

A4 -669 365.0 -1 684 915.0 -563 713.9

valid T [°C]

ref

0-300 0-250 0-225 0-225 0-250 0-250 0-160

10 15 8 8 15 15 10

Model Implementation. To apply the model in practical calculation, the upper part of the column between the gas exit temperature and gas temperature at the start of condensation is modeled as n discrete elements represented by equal temperature intervals. The simulation algorithm for the entire column is sketched in Figure 5 for the case of known raw gas NH3 content. The CO2 concentration in clean gas must be known. Iteration starts at the gas exit with an estimated value for the NH3 concentration and proceeds downward with increasing temperature until the raw gas composition is reached at the section where condensation begins. The gas-liquid equilibrium requires the solution of a nonlinear system of eight equations (eqs 17-24) in each element. To improve the convergence behavior of the commercial solver used, the system can be reduced to a one-parameter problem with one concentration, e.g., mOH-, as the variable. Therefore, eqs 17-23 are transformed to a logarithmic scale, presenting a 7 × 7 linear

n

∑ mi(total)(Tk) ∆YCOND,k k)1

(28)

It is important to notice that also the gas-phase composition may change along the column’s height influencing mi(total)(T). For the gas phase, the partial mass balance of the electrolyte components NH3 and CO2 in the temperature interval ∆T can be written as

∆Yi,k ) mi(total)(Tk) MH2O∆YCOND,k

(29)

with the molar mass of water MH2O in kilograms per mole. The molar load Yi of the electrolyte in the dry gas is related to the molar fraction yi used in eqs 17 and 18 according to

Yi ) yi(1 + YH2O)

(30)

Equations 29 and 30 imply that the impact of the concentration changes of NH3 and CO2 on the total amount of dry gas is negligible because of the small quantities of NH3 and CO2 removed specific to the total dry gas flow. The same applies for the aqueous liquid phase, where the low concentrations of dissolved species do not affect the total liquid mass flow rate. The model presented above allows the calculation of the gas-phase NH3 profile along the temperature coordinate if the clean gas CO2 concentration and at least one of the NH3 concentrations in condensate, raw gas, or clean gas are known. The main dry gas components, the raw gas water content, and the in/out temperatures of the gas and liquid streams are also necessary to determine the energetic section. The required raw gas water content YH2O,in can be substituted by the total specific amount of condensate YCOND according to eq 12.

Figure 5. Calculation algorithm for a given input NH3 concentration Y′NH3,in.


Figure 6. Specific water load in gas and specific amount of condensate versus the gas-phase temperature inside the countercurrent column.

Figure 7. Partial pressures of CO2 and NH3 for different NH3 concentrations in raw gas versus the gas-phase temperature inside the column.

system that can be solved straightforward and yielding a set of concentrations for a given mOH-. The solver finds the mOH- for which the electric balance (24) is fulfilled. When the algorithm has reached the section where condensation starts, the calculated raw gas NH3 concentration is compared to the target value and the solver generates a new starting guess for the exit gas concentration until the values match. A minor modification in the algorithm allows the calculation of both gas-phase NH3 compositions if the total amount of NH3 in the condensate mNH3(total),out is known. The algorithm is fast and robust with respect to convergence. Numerical errors fade with an increasing number of elements and can be considered insignificant for n g 20. 4. Results and Discussion Concentration Profiles in the Scrubber. As introduced above, the gas-phase temperature is used as a coordinate along the column’s height. In typical operation conditions of the Guessing scrubber, concentration profiles are presented for a producer gas composition according to Table 1, gas in/out temperatures of 170/50 °C, solvent in/out temperatures of 40/80 °C, and a water content in the raw gas of 40 vol %. The raw gas NH3 content is varied from 500 to 2000 vppm (dry). Figure 6 shows the water load in the gas stream and the specific amount of condensate versus gas-phase temperature. According to the model assumption that no re-evaporation of condensate occurs in the lower part of the column, the water loads are constant in that region. The partial pressures of NH3 and CO2 in the gas phase are presented in Figure 7. The CO2 pressure is increasing toward the top of the column because the absorbed CO2 does not significantly affect the gas composition and the total amount of gas is decreasing with condensation. The same effect results in almost unchanged partial pressures of NH3 along the column’s height. The actual NH3 removal in the scrubber is shown by NH3 in dry gas in Figure 8. The separation efficiency defined as

ηsep,i ) 1 -

Yi,out Yi,in

(31)

is 50% for 500 vppm of NH3 in dry raw gas and only 30% for 2000 vppm of NH3 in dry raw gas.

Figure 8. NH3 load on dry gas for different raw gas NH3 concentrations versus the gas-phase temperature inside the column.

Influence of Process Parameters on NH3 Removal. According to the model, the NH3 separation efficiency of the scrubber depends on the following process parameters: (i) clean gas temperature; (ii) raw gas water content; (iii) CO2 content in gas; (iv) NH3 content in gas. The clean gas temperature and raw gas water content determine the amount of condensate according to Figure 9. A comparison between Figures 9 and 10 shows that the effect of TG,out and yH2O,in on the NH3 separation efficiency cannot be reduced to the parameter YCOND only. The reason is the nonlinearity of the gas-liquid equilibrium. The effect of the clean gas CO2 content and raw gas NH3 content on the separation efficiency is shown for a fixed clean gas temperature and raw gas water content in Figure 11. The corresponding pH value of the condensate is presented in Figure 12. Comparison with Measured Data. Extensive NH3 measurements have been carried out at the Guessing plant. The solution of partial energy balances on the column allows the calculation of the liquid-phase temperature corresponding to gas-phase temperatures. For the several sampling runs, the calculated pinch temperatures ∆TGL,C at the section where condensation starts vary between 0.5 and 2.0 K. If the measured raw

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Figure 9. Specific amount of condensate versus the raw gas water content and clean gas temperature. Figure 12. pH value of the condensate at 298.15 K versus the clean gas CO2 content and raw gas NH3 content.

Figure 10. NH3 separation efficiency of the scrubber versus the clean gas temperature and raw gas water content. Figure 13. Comparison of the clean gas NH3 concentration measured with the value calculated from the measured raw gas concentration.

Figure 11. NH3 separation efficiency of the scrubber versus the clean gas CO2 content and raw gas NH3 content.

gas concentration is taken as the input parameter for calculation, the clean gas concentration and the separation efficiency respectively can be calculated. Figure 13 shows measured clean gas NH3 values dependent on the values calculated from the corresponding raw gas NH3 concentrations and the specific operation parameters

such as temperatures, raw gas water content, and CO2 content in clean gas. The accuracy of the analytical determinations of gas-phase NH3 is estimated to be about (100 vppm (dry) for NH3 concentrations in the range of 300-2000 vppm (dry). The same error must be considered for the calculated values, which, of course, depend on the measured raw gas concentrations. The separation efficiencies, calculated from the gas concentration before and after the scrubber, are compared in Figure 14. The model tends to overestimate the NH3 removal effect of the scrubber, which results in slightly lower calculated clean gas NH3 concentrations than those found by the measurements. Possible reasons for the deviations between the model and measurements could be the model assumptions and measurement inaccuracies. One possible explanation is that, despite the model assumption, NH3 is transferred through the organic phase from the hot water phase back to the gas phase. This may happen in the hot lower part of the column and also in the equalization tank. A global mass-transfer parameter for NH3 could be introduced to account for these deviations. To avoid the


the prediction of clean gas values even though a tendency toward an overestimation of NH3 removal can be observed. If gas volatilization from the emulsified water is the reason for the lower NH3 separation efficiencies measured, the calculation results represent an upper bound for NH3 removal in the scrubber. Further investigation will focus on the variation of operating parameters in the Guessing gas-cleaning system in order to learn more about the possibilities of increased NH3 removal by simple measures. Literature Cited

Figure 14. Comparison of the measured and calculated NH3 separation efficiencies of the scrubber.

measurement errors, all triples of NH3 concentrations (raw gas, clean gas, and condensate) that did not fulfill the mass balance were excluded from consideration. 5. Conclusions Within the present work, a relatively simple model of coabsorption of NH3 and CO2 into condensing water in a countercurrent organic solvent scrubber has been developed based on the following assumptions: (i) exiting gas relative humidity of 100%; (ii) no radial gradients and no axial mixing in the column; (iii) pure organic solvent at the top of the column; (iv) no mixing between water and the organic phase; (v) no solubility of NH3 in the organic solvent; (vi) gas-liquid equilibrium during condensation; (vii) condensate behaving as an ideal electrolyte solution; (viii) no re-evaporation of the condensate in the hot lower part of the column; (ix) no gas volatilization from emulsified water. The gas-phase temperature is applied as a height coordinate of the column. Therefore, no geometry data are needed in the modeling, and the model does not contain any empirical parameter that needs to be determined from measurements. An algorithm is presented that allows the calculation of clean gas and condensate NH3 contents if NH3 in raw gas and certain process parameters are known. Calculations show for the typical operation range of the producer gas scrubber that the partial pressure of NH3 does not change significantly along the column’s height because of a decrease in the total gas quantity (condensation of water). The separation efficiency is strongly dependent on the gas exit temperature, raw gas water content, CO2 content in gas, and raw gas NH3 content. While tar separation is very efficient, the NH3 separation efficiencies in the possible operation range of the scrubber are limited to about 50%. Quantitative removal of NH3 from the producer gas can be reached with further gascleaning units, which may be economically critical at small plant capacities. On the other hand, primary catalytic measures in the hot raw gas can be set in order to decrease the load of the raw gas and to reach certain NH3 limits in clean gas after the tar scrubber. Extensive NH3 measurements have been performed at the organic solvent scrubber in Guessing, Austria. The model evaluation shows satisfactory agreement for

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Report TAE 804; Technion Israel Institute of Technology: Technion City, Haifa, Israel, 1997 (http://garfield.chem.elte.hu/Burcat/ burcat.html).

Received for review April 23, 2004 Revised manuscript received December 7, 2004 Accepted December 10, 2004 IE049669V

Removal of NH3 from Biomass Gasification Producer Gas by Water

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