Ind. Eng. Chem. Res. 2003, 42, 6755-6764
6755
Experimental Fact-Finding in CFB Biomass Gasification for ECN’s 500 kWth Pilot Plant S. R. A. Kersten,† W. Prins,‡ A. van der Drift,† and W. P. M. van Swaaij*,‡ Netherlands Energy Research Foundation (ECN), P.O. Box 1, 1755 ZG Petten, The Netherlands, and Faculty of Science and Technology, University of Twente, P.O. Box 217, 7500 AE Enschede, The Netherlands
CFB biomass gasification has been studied by experimentation with ECN’s pilot facility and a cold-flow model of this plant. Data obtained by normal operation of this plant and the results of some special experiments have provided new insight into the behavior of circulating fluidized bed reactors and CFB biomass gasifiers in particular. It has been found, for instance, that char gasification reactions and reform reactions of lower hydrocarbons do not proceed in an atmospheric CFB. Moreover, it appears that the axial and radial mixing of gas and solids is not ideal and that the coupling between the prevailing solids circulation rate, the gas velocity (profile), and the solids hold-up does not provide optimal reactor conditions. As a consequence, problems are observed in practice, such as insufficient thermal efficiency and an excessive concentration of contaminants (e.g., tars) in the product gas of CFB gasifiers. A two-dimensional engineering reactor model has been developed, partially on the basis of the results obtained from measurements with the pilot plant and the cold-flow model. The predictions of this model are in good agreement with the observed overall performance of the plant, as well as the measured axial and radial gas concentration profiles inside the riser. A novel multistage fluidized bed reactor concept has been tested (at laboratory and pilot scale) to investigate whether the problems encountered in conventional CFB gasifiers could be resolved. In this novel reactor, a separate char combustion zone is created that results in a significant increase of the thermal efficiency. 1. Introduction In many western countries, biomass is supposed to play an important role in the future energy production system. Circulating fluidized bed (CFB) biomass gasification is one of the possible conversion methods available to allow the channeling of biomass into modern energy conversion equipment. CFBs have always been considered attractive, because of their high throughput per unit volume reactor, high mass and energy exchange rates, and simple scale-up. Papers concerning the interpretation of data from pilot-scale CFB reactors, in which a real chemical process is carried out, are scarce in the open literature.7,12,18,33 The majority of the research concerning riser systems is still focused on hydrodynamics. At the same time, most papers concerned with chemical reactions in CFBs do not consider the physical processes involved. They primarily discuss the overall performance of the reactor, e.g., the influence of the process conditions on the conversion and selectivity. In this paper, the performance of CFB biomass gasifiers is discussed in relation to experimental data from ECN’s pilot gasifier (ca. 500 kWth) and a corresponding cold-flow model. Numerous input-output data sets of regular tests have been collected and are used to characterize the CFB biomass gasifier. Apart from these overall performance data, the results of special experiments such as the measurement of axial and radial gas concentration profiles, methane injection, and bed material (char) analysis, are also presented. The * To whom correspondence should be addressed. E-mail:
[email protected]. † Netherlands Energy Research Foundation (ECN). ‡ University of Twente.
cold-flow model of the pilot plant is used to examine gas mixing and to visualize the spatial distribution of biomass. All data obtained are analyzed to define the phenomena that should be incorporated into an engineering reactor model for the model to be successful. Predictions of such an engineering model are validated against the experimental results. On the basis of the insight gained through experimentation with the existing CFB pilot plant, a new multistage reactor concept is designed specifically to overcome the problems encountered in conventional systems.20 The performance of this novel reactor is discussed briefly as well. Biomass gasification is a conversion process in which solid biomass is transformed to a gaseous energy carrier (CO, CO2, CH4, C2+, H2, H2O). The reaction steps involved are as follows: (i) heating and thermal decomposition (pyrolysis) of the biomass, in which the biomass is converted into three lumped component classes, viz., permanent gas, char (a solid with a typical composition of CH0.2O0.1), and primary tars; (ii) cracking (thermally and/or catalytically) of the primary tars to permanent gases and secondary tars and further cracking of these secondary tars to ternary tars and soot; (iii) heterogeneous gasification reactions of the char produced by biomass pyrolysis and homogeneous gas-phase reactions (e.g., the shift reaction); and (iv) combustion of the char and oxidation of combustible vapors. In common practice, the ratio of air (oxygen) to fuel is defined by the equivalence ratio (λ), which is the ratio between the actual amount of oxygen supplied to the system and the amount of oxygen needed for complete combustion of the fuel. The equivalence ratio required to reach and maintain a certain reactor (gasification) temperature is primarily dependent on the moisture content of the fuel, the amount (and temperature) of
10.1021/ie020818o CCC: $25.00 © 2003 American Chemical Society Published on Web 11/15/2003
6756 Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003 Table 1. Range of Operating Conditions and Typical Gas Composition Operating Conditions biomass feed rate moisture content feed air feed rate steam feed rate temperature gas velocitya sand flux
2. Experimental Set-ups and Measurement Techniques (a) Pilot-Scale Gasifier at ECN. The circulating fluidized bed (CFB) gasifier under consideration in this study is called BIVKIN and is owned by the Netherlands Energy Research Foundation (ECN) in Petten, The Netherlands. It is an atmospheric air-blown facility suitable for a biomass throughput of up to 130 kg/h (600 kWth). Air can be fed to the riser through the primary nozzles at the bottom (primary air) or at several other axial positions (secondary air). The plant is equipped with various feeding systems that can be used simultaneously to allow the feeding of fuel mixtures. The fuel is introduced just above the primary air nozzles, or at a height of 1.0 m above the primary nozzles. The riser is 20 cm in diameter and 6 m high. Further included in the circulation loop are a cyclone, a downcomer, and a seal bubbling fluidized bed. The bed material used is sand that is 0.4-0.6 mm in diameter (Dsauter ) 0.53 mm). After the first cyclone, the gas is led through a second cyclone; dust and ashes are collected after this cyclone. Figure 1 shows a scheme of the facility; the range of operating conditions and a typical gas composition are given in Table 1. In the higher part of the riser, above the acceleration zone, the pressure drop is typically 100-200 Pa/m..7 During each gasification experiment, the concentrations of O2, H2, CO, CH4, and CO2 in the product gas were measured continuously using paramagnetic (O2), TCD (H2), and IR detectors, respectively. The quantities of nitrogen, ethene, acetylene, ethane, benzene, toluene, and xylene were measured discontinuously (every 10 or 20 min) by gas chromatography. The sample point is
range
kg/h wt % Nm3/h kg/h °C m/s kg/(m2‚s)
60-130 7-30 80-110 0-35 750-900 6-9 10-25
Gas Composition (vol % dry)b CO CO2 CH4 C2H4 C2H6 C6 + C7c H2 H2Od N2
Figure 1. Scheme of CFB gasification facility at ECN.
steam added, and the heat loss to the environment. Because the absolute heat loss of a pilot-plant gasifier to the environment is essentially constant at a fixed reactor temperature, varying the biomass feed rate changes the relative heat loss. In this paper, the relative heat loss (dimensionless) is defined as the absolute heat loss to the environment divided by the product of the fuel feed rate and the lower heating value (LHV) of this fuel. By increasing the fuel feed rates to high values, predictions can be made about operation at large scale (low relative heat loss). A large-scale gasifier will, in general, have a relative heat loss of less than 1%. At present, the thermal efficiency (too low) and the content of tar and dust particles in the product gas (too high) are considered to be the most important performance indicators (i.e., problems) in CFB biomass gasification technology.
units
14.2 16.4 4.0 1.37 0.08 0.03 11.9 12.4 51.4
a Superficial velocity based on the flow at the outlet of the riser. Operating conditions of this specific experiment given in Table 2. c Including benzene, toluene, and xylenes. d Volume percent wet.
b
located after the second cyclone. Occasionally, the water content in the product gas was determined by condensation and absorption methods. Two methods were applied to determine the tar content of the product gas: (i) “light” tars were measured by the SPA method5 (these light tars comprise molecules from xylene to about fiveringed compounds) and (ii) “heavy” tars were measured gravimetrically. At five axial positions, specially developed probes were installed to withdraw gas samples from the interior of the riser. These probes could be moved freely over the reactor diameter, so that full radial profiles could be obtained at each axial position.18 The same probes were also used to inject tracer gases. At nine axial positions, the bed temperature and the pressure were measured. To gather information on the amount and composition of the solids in the system () riser + downcomer + seal fluidized bed) under real gasification conditions, the following experimental procedure was applied: (i) The gasification experiment was stopped abruptly by switching off both the biomass and the air supply systems. (ii) Then, the system was flushed with nitrogen until the temperature was low enough (during this period, there was no solids circulation) that the solid material could be handled. (iii) Finally, a solids sample was collected from the system. Alternatively, the amount of carbonaceous material present in the system was measured by char burn-out experiments. In that case, only the biomass feed was stopped while the air flow through the system was maintained. Through measurements of the CO2 and CO fractions in the product gas, the carbon hold-up could be calculated. Many experiments with approximately 15 different fuels (wood, manure, grass, ties, chicken sludge, etc.) have been carried out in the ECN pilot facility.7 It has been shown that a CFB is suitable for the gasification of different kinds of biomass of varying composition. The closures of the mass, atom, and energy balances were usually very good, typically within 5% accuracy. Table 2 gives an example of the mass and energy balances of a regular experiment. The performance of ECN’s pilot CFB gasifier in terms of thermal efficiency, carbon conversion, and tar concentration in the product gas appears to be typical for
Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003 6757 Table 2. Mass and Energy Balances of a Typical Pilot-Plant Experimenta Input temperature total mass mass C mass H mass O mass N mass ash mass restb LHV SH Hvap
units
biomass
air
N2
°C kg/h kg/h kg/h kg/h kg/h kg/h kg/h MJ/h MJ/h MJ/h
25 92 40.27 5.90 43.35 0.34 2.05 0.08 1490.3 5.1 24.3
270 131.8
25 4.43
30.7 101.1
4.43
total 228.2 40.27 5.90 74.05 105.87 2.05 0.08 1490.3 37.6 24.3
32.5
Output temperature volume flowd total mass mass C mass H mass O mass N mass ash LHV SH Hvap
units
gas
ashes + dustc
tar
°C Nm3/h kg/h kg/h kg/h kg/h kg/h kg/h MJ/h MJ/h MJ/h
851 164.4 215.95 34.69 5.57 71.59 104.1
851
851
5.36 3.60
1.57 1.27 0.10 0.16 0.05
994 248.2
1.76 93.8 3.2
total
47.6 1.9
222.88 39.54 5.67 71.75 104.15 1.76 1135.4 253.3
Balance Closure units masse
total mass C Xgas (excl tars) Xgas (incl tars) Xashes/dust mass H mass O mass N mass ash energy (LHV)f
% % % % % % % % % %
97.7 98.2 86.1 89.3 91.1 96.1 96.9 98.4 85.9 97.0
λ ) 0.275, 10 wt % moisture in the feed, no steam added, heat loss ) 20 kW; for the gas composition of this particular experiment, see Table 1. b For example, Cl, S, and metals. c Both components contain ashes (minerals) and carbon (char). d Dry gas. e % ) (out/ in) × 100. f % ) (LHV + SH)out/[(LHV + SH - Hvap)in - Q] × 100. a
CFB gasification in general. This was concluded by comparison with data published for other CFB biomass gasifiers (i.e., 18,29 45,25 20,4 10,24 27,16 and 1001 MWth). This comparison showed that the values of the main performance indicators (thermal efficiency, carbon conversion, tars in the product gas) were similar and, more importantly, did not dependent on the capacity. For the concentration of tar in the product gas, values of 1-10 g/Nm3 have been reported (ECN, 8 g/Nm3). The reported cold-gas efficiencies vary from 55 to 75% (ECN, 62%), and the carbon conversion values are typically in the range of 83-95% (ECN, 90%). An expected influence of increasing throughput (increasing reactor size) is that the thermal efficiency would increase as a result of the decreased relative heat loss. (b) Data Interpretation. Each type of fuel used and the collected char, tar, and dust fractions were subjected to proximate and ultimate analyses, as well as measurement of the heating values, all according to the standards.15 In the mass balance, time-averaged values of the concentrations measured during continuous operation were used. This is a reasonable assumption, given that it has been found that the product-gas concentrations do not fluctuate much in time (usually
σi/xi < 0.05) during continuous operation under fixed gasification conditions. The process temperature was chosen as the mean of the axial bed temperatures, because the maximum temperature difference between the bottom and top of the riser was always less than 30 °C under normal operating conditions. Two methods were applied to calculate the degree of carbon conversion: (i) dividing the amount of carbon in the gas phase (including or excluding tars) by the amount of carbon in the fuel (Xc,gas) and (ii) dividing the amount of carbon in the ashes and dust collected in the second cyclone and by the filters by the amount of carbon in the fuel (1 - Xc,ashes/dust). The difference between the two calculated conversions was used, in conjunction with the H, O, N, and energy balances, to assess whether a certain experiment had sufficient mass/energy balance closure to be regarded as successful. (c) Auxiliary Cold-Flow Modeling. The cold-flow setup is a downscaled version of the pilot-scale gasifier. Its linear dimensions are approximately one-half of the original ones, and the setup is made largely from glass. This cold-flow setup enables the operator to actually see what is happening during operation. In this cold-flow model, occasional adjustments and modifications can be tested in an easier and less expensive way. At 15 equidistant axial positions, pressure tabs are present, so that pressure profiles can be recorded. The solids flux was measured by a valve in the downcomer that bypasses the solids stream into a collector vessel. Air was used as the fluidization agent. In the cold-flow model, radial gas mixing can be investigated by performing tracer-gas (He, CO2) experiments. Moreover, adding batches of such particles to the cold-flow riser simulated the spatial biomass and char distribution under gasification conditions. 3. Experimental Fact-Finding Extensive experiments were carried out to determine the effects of various parameters on the gasifier performance. In this section, the observations/conclusions are collected, derived from regular tests as well as from special experiments to characterize the reactor system in more detail. In this paper, the results are presented in summarized form. For a more detailed description of the experimental methods employed, the reader is referred to refs 7 and 18-20. (a) Effect of Particle Size. No significant effect of the fuel particle size (particles with dimensions Lp × dp ranging from 2 mm × 2 mm to 40 mm × 10 mm were used) on the product-gas composition, carbon conversion, and tar content in the product gas has been observed.8 Under the prevailing conditions, external mass transfer limits char combustion. As a consequence, the combustion rate of small particles is much larger than that of coarse particles, especially on the basis of the particle’s volume. However, analysis of the bed material (see below) showed that, even when large biomass particles are fed to the reactor, there are still many small char particles present in the riser. Moreover, it has been found that only the smallest char particles are combusted (see below). The combination of these two observations might explain why the particle size of the feed hardly influences the carbon conversion. (b) Gasification of Char. From the results of a separate pilot-plant test with steam/air and wood char as the feedstock, it was concluded that carbon gasification reactions do not proceed to a significant extent
6758 Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003
Figure 2. Data from ECN’s pilot plant. The carbon conversion (excluding tars) and the CO2, CO, and H2 yields are plotted versus the H/C ratio of the feed (including biomass moisture and steam). Fuel type: demolition wood. In all experiments, the temperature was about 850 °C. The moisture content of the biomass varied from 9 to 11%, and the equivalence ratio varied from 0.25 to 0.32. The H/C ratio was primarily varied by addition of steam. Lines in the graph represent trends.
under atmospheric biomass gasification conditions in a CFB (700-1000 °C). It can be deduced that nearly all of the carbon in the gas phase can be ascribed to oxidation reactions.19 The same conclusion was drawn from a series of biomass gasification experiments during which steam (100-120 °C) was added at otherwise comparable conditions (λ, T). It was found that the addition of steam leaves the carbon conversion unaffected and results in increased CO2 and H2 yields and a decreased CO yield, indicating that the addition of steam influences only the shift equilibrium (see Figure 2). (c) Light Hydrocarbon Yields. Experimental results show that the yield of small hydrocarbons is not notably affected by the operating conditions in the following ranges: T ) 740-910 °C, λ ) 0.17-0.40, and H/C ) 1.6-2.8. The highest value of H/C corresponds to a steam-to-biomass ratio of about 0.30 kg/kg. The carbon based yields (Yc) are CH4 ) 8%, C2H4 ) 5%, and C6H6 ) 2.5%. These yields remain constant when significant amounts of steam are added to the reactor, which indicates that steam-reforming reactions do not proceed. To further probe the previous observation, some dedicated experiments were performed. During pilotplant runs, a known amount of methane was injected continuously into the riser at a position 1 m above the primary air nozzles where oxygen was never detected. Extra CH4 was injected for 45 min. Through a comparison of the CH4 yield before and after this additional injection, it was found that the additional methane did not reform in atmospheric noncatalytic CFB biomass gasifiers operated at temperatures below 1000 °C. From the experimental data,17 it was already clear that CH4 and C2H4, as well as C6H6, do not crack thermally in the presence of steam and hydrogen up to temperatures of 1300 and 1050 °C, respectively, at typical gasification residence times. Now, it has been made clear as well that the char and minerals present in CFB gasifiers show no catalytic activity for the reforming of CH4. (d) Extrapolation of Pilot-Plant Results. In contrast to a large-scale CFB plant, a small pilot plant
Figure 3. Data from ECN’s pilot plant. The equivalence ratio, carbon conversion (excluding tars), LHV (wet gas) of the product gas, and cold-gas efficiency are plotted versus the biomass feed rate. Fuel types: willow, beech, and demolition wood. In all experiments, the temperature was maintained at approximately 850 °C, and the moisture content of the biomass varied between 9 and 15%. No steam was added. A feed rate of 60 kg/h biomass corresponds to approximately 13% relative heat loss, 130 kg/h to 6% relative heat loss. Lines in the graphs represent trends.
always experiences significant heat loss per kilogram of biomass fed to the system. This offers the opportunity to operate a pilot plant over a wide range of equivalence ratios at constant temperature by varying the biomass feed rate, viz., from λ ) 0.17 to λ ) 0.4 at 850 °C for ECN’s CFB. In this range, when the feed rate is varied, the equivalence ratio can always be changed in such a way that a corresponding change in the relative heat loss compensates the effect in the overall heat balance. From these experiments, it was concluded that the thermal efficiency of the ECN pilot facility is essentially constant over a wide range of applied equivalence ratios at a fixed temperature (see Figure 3). The thermal efficiency is defined as
ηLHV )
gas gas Θproduct (λ,T)LHVproduct (λ,T) v v biomass biomass Θm,daf LHVm,daf
)
gas (λ,T) Rgas(λ,T)LHVproduct v biomass LHVm,daf
In this equation, Rgas is the amount of gas produced per kilogram of feed (Nm3/kg). The nearly constant thermal efficiency (ηLHV, Figure 3) can be explained by the following line of reasoning: (i) When the relative heat loss decreases, the equivalence ratio has to be decreased to maintain a constant operating temperature (see Figure 3). (ii) As a result, the LHV of the product gas increases (see Figure 3) because there is less combustion and dilution with inert N2. (iii) A lower equivalence ratio results in a lower carbon conversion (see Figure 3). (iv) Because of the lower carbon conversion, the amount of
Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003 6759
gas produced per kilogram of feed decreases. (v) Apparently, the changes in the LHV of the product gas and in the gas production (Rgas) compensate each other roughly when the equivalence ratio is varied at constant temperature. It can be concluded that, for the optimization of a conventional air-blown CFB biomass gasifier at a fixed operating temperature, a choice remains between a high LHV (low λ) of the product gas and a high carbon conversion (high λ). For a large-scale process, this might indicate that a high carbon conversion can be attained only by operating at higher temperatures or by building heat sinks into the gasifier (steam generation). However, especially considering the high alkali metal concentration of biomass feedstock, at higher operating temperatures (say, 950 °C), measures have to be taken to avoid solids agglomeration. (e) Carbon Balance. In the gasification process, char is formed primarily by the pyrolysis step and, to a much smaller extent, by tar polymerization reactions. Char can be converted by two mechanisms, viz., by heterogeneous gasification reactions and by combustion. During continuous operation, carbon (char) leaves a CFB gasifier via the cyclone(s), which have a limited separation efficiency. Particles collected in the cyclone of ECN’s gasifier are smaller than 0.3 mm and contain typically 60 wt % carbon and 30 wt % ashes. As a comparison, the carbonaceous particles inside the riser and seal fluidized bed contain ca. 80 wt % carbon and 12 wt % ashes. Char gasification reactions do not contribute to the carbon conversion to a significant extent under atmospheric biomass gasification conditions in a CFB (see above), which leaves combustion as the dominant mechanism to convert char to the gas phase. From the measured axial gas concentration profiles, it can be concluded that devolatilization of the fuel, tar cracking reactions, and combustion reactions proceed concurrently in the bottom region of a conventional CFB riser (see below). In the bottom zone of the riser, oxygen is present next to char and combustible gases. Therefore, in practice, the amount of air theoretically required per unit biomass for complete combustion of the pyrolyis char will not be sufficient, because other gaseous combustibles are competing in the consumption of oxygen. To verify that there is indeed competition between char and volatiles combustion reactions in ECN’s CFB gasifier, the measured carbon conversion was plotted versus the applied equivalence ratio. The mass balance lines for the hypothetical cases that oxygen reacts only with char to form CO or CO2 are presented as well in Figure 4. The carbon conversion at an equivalence ratio of zero (pyrolysis) at the prevailing conditions was estimated from the compositional analysis of the carbonaceous material found in the riser combined with the composition of the feed material. From Figure 4, it becomes clear that, in a CFB, oxygen reacts not only with char. The measured carbon conversions data points are even below the C + O2 f CO2 line, while it has been observed through char burn-out tests and tests with char as the feedstock that CO is also formed during char combustion at λ e 1.19 From these char burn-out tests, it can also be concluded that, in ECN’s pilot gasifier, the carbon hold-up is, in principle, high enough to convert all incoming oxygen under standard gasification conditions. In fact, under normal operating conditions, there was ca. 4-6 kg of char was present in the system, which allowed for 10-15 min of
Figure 4. Carbon conversion (including tars) versus the equivalence ratio. Experimental data points are presented next to the mass balance lines for the hypothetical cases that oxygen reacts only with char to form CO or CO2. Conditions: see Figure 3.
combustion-mode operation before oxygen broke through. During gasification, oxygen is absent in the largest part of the CFB’s volume, leaving attrition/fragmentation as the only mechanism to reduce the char particles in size. However, also in the oxygen-rich zone, the attrition rate is higher than the combustion rate for char particles larger than 2 mm.27 Analyses have shown that there is no variation in the composition of the carbonaceous bed material as a function of its particle size (in the range of dp ) 0.4-10 mm, carbon ≈ 80 wt % and ashes ≈ 12 wt %). From the above, it can be concluded that the small char particles are mostly the remains of larger particles after mechanical splitting and attrition and that only the combustion of the smallest char particles contributes significantly to the carbon conversion to the gas phase. (f) Spatial Distribution of Biomass in the Riser. It has been observed that large biomass particles are present in cold-flow risers (Lp ) 1-3 cm, dp ) 1 cm, F ) 720 kg/m3) in the top of the riser already at moderate solids fluxes [10-30 kg/(m2‚s)]. The fact that large biomass particles can rapidly move to the top of the riser might contribute to the amount of tar that is measured in the product gas of CFB gasifiers, because the remaining residence time of the tar that evolves from these particles is not sufficient for complete conversion. Large (Lp > 8 mm) char particles were also found in the seal fluidized bed, indicating that biomass might even be ejected from the riser before being completely gasified. (g) Axial Temperature Gradient. Once a certain flux of inert solids has been achieved, the axial temperature gradients in the riser of a CFB biomass gasifier are small and limited to at most 30 °C from bottom to top. A ratio of the sand circulation flux to the biomass feed flux of about 25 is enough to keep the temperature gradient within this limit. (h) Radial and Axial Concentration Profiles. Figure 5 presents a typical example of centerline longitudinal gas fraction profiles in the pilot riser measured during a standard gasification test (demolition wood, dp ) 10 mm, Lp ) 10-30 mm, F ) 1200 kg/ m3). For other types of wood with a lower density (F ) 720 kg/m3), similar profiles were found. From these profiles, even though only the centerline volume fractions are presented in Figure 5, it can be concluded that most of the gas is produced in the first 1.5 m above the
6760 Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003
Figure 5. Measured and predicted axial volume fraction profiles for CO2, CO, H2, and CH4 at the center of the reactor. λ ) 0.275. Other conditions are listed in Table 2. The model details are given in Table 3. The axial coordinate value 0 corresponds to the (end) position of the primary air nozzles.
Figure 6. Measured and predicted radial volume fraction profiles for CO2, CO, H2, and CH4 at a reactor length of 2.75 m. λ ) 0.275. Other conditions are listed in Table 2. The model details are given in Table 3.
primary air nozzles, a distance representing 25% of the total length. This indicates that the majority of all three subprocesses of gasification (devolatilization, tar cracking and combustion) takes place in the bottom part of the riser (say, 20% of the total length). At the lowest sample point (0.7 m above the primary air nozzles), oxygen was never detected in runs with only primary air, which indicates that the oxidation reactions are very fast; complete oxygen conversion in less than 0.1-s gas residence time is achieved. Measured radial gas fraction profiles for the main components are plotted in Figures 6 and 7. While referring to these figures, it should be noted that the predicted lines will be discussed further in section 4. Profiles of CO, CO2, CH4, and benzene (not shown here) exhibit parabolic behavior. Fractions near the wall can be up to 3 times higher than fractions in the center. Radial profiles such as those in Figure 6 are already observed at a height of 1 m above the primary air nozzles. The observed radial distributions of gas components can be caused by various phenomena: (i) One potential cause is a nonuniform radial distribution of biomass. More biomass near the wall will induce more gas production near the wall. This results in higher product-
Figure 7. Measured and predicted radial volume fraction profiles for H2 at reactor lengths of 2.75, 4.00, and 4.75 m. λ ) 0.275. Other conditions are listed in Table 2. The model details are given in Table 3.
gas component fractions near the wall if radial gas mixing is insufficient. Nonuniform radial solids distributions in risers have been reported.28 (ii) Another is a parabolic gas velocity profile. As a result of such a velocity profile, a radial gas residence time distribution would arise. A longer residence time near the wall, corresponding to a lower gas velocity, would lead to higher gas fractions assuming that the volumetric gas production rate is uniform over the radius, because there would be less dilution with nitrogen. Parabolic gas velocity profiles have been observed in comparable reactor configurations by a number of authors.6 (iii) The observed radial distributions might also result from heterogeneous secondary tar cracking. If the cracking rate of secondary tar were proportional to the solids hold-up (char), more cracking would occur near the wall, because in a riser, more solids are typically present near the wall. (iv) Combinations of the three reasons mentioned above might also lead to the observed distributions. Information concerning the radial mixing in the riser is obtained from Figure 7, in which radial H2 profiles at three different axial positions are shown. If it is true that only little gas production takes place after 2 m of the reactor length, radial mixing must be poor. On the basis of the measured radial gas concentration profiles and the standard dispersion model used to interpret the results, it can be concluded that the average radial Pe´clet number of the dilute region (L ) 1.25-6 m) is on the order of 1000 or higher, which corresponds to a radial dispersion coefficient of about 5 × 10-4 m2/s. Average radial Pe´clet numbers are assumed to be invariant in relation to the spatial coordinates. Also, from tracer-gas measurements in the cold-flow riser, it can be concluded that radial gas mixing in risers is poor.18,19 Experiments showed that the radial CO2 profile is almost flat during gasification runs in which air is supplied only through the bottom nozzles (see Figure 6). However, during runs with secondary air (not shown here) and during biomass combustion experiments (see Figure 8), strong radial CO2 profiles (differences) were measured. This would indicate that the existence and the shape of the radial CO2 profile is related to the local oxygen-to-fuel ratio. In CFB gasification, most of the CO2 is formed by fast reactions of oxygen with carbon
Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003 6761
that a local excess of oxygen in this region can cause deviations from the flat CO2 profile. During combustion runs, the local excess in oxygen instead of fuel and the low radial gas exchange rates probably cause the observed radial CO2 and O2 concentration profiles (see Figure 8). 4. Modeling
Figure 8. Measured and predicted axial volume fraction profiles of CO2 and O2 in the center and near the wall of the riser during a biomass combustion experiment. λ ≈ 2.5, Θbiomass ) 12 kg/h, m 3 Θair v ) 125 Nm /h, T ) 840 °C. The model details are given in Table 3.
(char), CO, and hydrocarbons. Obviously, there is no radial O2 profile just after the primary air nozzles. Locally, there is always an excess of combustibles when air is supplied only through these primary air nozzles. As a result, the radial concentration profile of CO2 will be relatively flat soon after the primary air nozzles, because the rate of the oxidation reactions is very high. All of the effects of the gas velocity profile and the voidage distribution are then of minor importance. If secondary air is injected, it is injected near the wall, so
For interpretation of the measured radial and axial gas concentration profiles, an engineering reactor model was developed. A frequently used reactor model for risers is based on the concept of a discontinuous coreannulus structure to account for radial effects.22 Unfortunately, the core-annulus model requires input data that are, in principle, unknown and difficult to assess (solids fraction near the wall, core-to-annulus mass-transfer coefficient). The model presented here is two-dimensional and based on empirical descriptions of the continuous radial profile of the axial gas velocity, radial gas dispersion, radial gas mixing, and radial solids segregation, without taking into account any shielding of particles in clusters. The empirical input for such a model has been studied frequently, both in theory and in practice, and the results of these studies are generally in reasonably good agreement with each other. It must be emphasized that the model used is an interpretation model, and therefore, it cannot be used for design and scale-up purposes. Recently, Wei et al.33 also applied such a model for the description of propylene ammoxidation in a riser. The modeling assumptions and input parameters of the base case are listed in Table 3. These input parameters
Table 3. Base-Case Model Parameters and Assumptionsa,b model parameter biomass char dense zone dilute zone radial gas mixing dense zone dilute zone axial distribution of solids biomass hold-up char hold-up in the riser location of biomass devolatilization (spatial biomass distribution) shape of the radial distribution of the axial gas velocity radial distribution of solids hold-up (sand + char + biomass) chemical kinetics
a
value(s)
comments
F ) 1200 kg/m3, Lp ) 2-3 cm, dp ) 1 cm F ) 300 kg/m3, dp distribution19 0-1.25 m 1.25-6 m
measured values
〈Dr〉 ) 0.003 m2/s 〈Dr〉 ) 0.0006 m2/s
average 〈Dr〉 values measured in the cold-flow setup and scaled to the hot conditions using Gaya´n et al.’s13 correlation relation of Li and Kwauk23 used based on measurements in the pilot plant
measured value based on measured pressure gradients in the cold-flow riser
average ) 〈〈0.97〉〉 vol % dense zone ) 〈〈4.2〉〉 vol % dilute zone ) 〈〈0.12〉〉 vol % estimated average ) 〈〈1.7〉〉 vol % 90% in the first 1.25 m above the primary air distributor 10% in the top 4.75 m of the riser R ) 5 in the relation of Godfroy et al.14
(1 - θR) ) γ 〈Uz〉 Uz
〈〉0.4 - (θ) 〈〉
0.4
- 〈〉
shape of profile based on measurements in the cold-flow CFB and predictions of a CFD model used to simulate pilot-plant riser21 according to the relation of Patience et al.26
) 4θ6
Wagenaar32 Wagenaar32 Van den Aarsen3 limited by external mass transfer; CO/CO2 ratio from Arthur2 Dryer and Glassman11 Dryer and Glassman11 same rate as for CH4 assumed very high rate assumed equilibrium assumed
pyrolysis cracking of primary tar cracking of secondary tar char oxidation CO oxidation CH4 oxidation C2H4 and C6H6 oxidation H2 oxidation shift
Details on the measurements can be found in Kersten.19
on average, 1 kg of char was found in the riser, which corresponds to 1.7 vol % Figure 5 shows that, for the prevailing conditions, most of the gas is produced in the bottom part of the riser
b
Experimental conditions listed in Table 2.
6762 Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003 Table 4. Model Predictions for Different Assumptions Regarding the Radial Axial Velocity and Solids Hold-up Distributions simulated cases model input
1
2
3
4
gas velocity profile radial biomass hold-up axial biomass hold-up radial char hold-up radial gas mixing
BCa
parabolic BC BC BC BC
flat BC BC BC BC
flat Ub U U BC
BC BC BC BC
predicted
product flow rate (kg/h)
measuredc
1
2
3
4
CO H2 CO2 CH4 C2H4 C6H6 H2O tar Xc,gas (incl tar) (%)
29.07 1.74 52.93 4.72 2.81 1.45 18.55 1.36 88.20d
27.61 1.76 55.98 6.09 3.36 1.62 17.59 0.041 89.89
26.57 1.73 56.24 5.95 3.34 1.61 18.03 0.038 88.62
30.48 1.85 55.26 6.19 3.39 1.64 16.43 0.056 92.80
33.76 1.94 54.31 6.27 3.43 1.66 15.54 0.058 95.94
a BC ) base case (kinetics included in the base case listed in Table 3). b U ) spatially uniform. c Measured values correspond to the experiment described in Table 2. d Excluding toluene and ethane.
of the present model were derived as far as possible from measurements in the pilot plant and in the cold-flow CFB (e.g., radial gas dispersion coefficient, char holdup, see Table 3). Other empirical input used in this model, such as the shape of the radial profile of the axial gas velocity and the axial and radial solids distribution, is taken from the literature (see Table 3). The chemical reactions included are also listed in Table 3. In Figures 5-8, the predictions of the reactor model are plotted together with the measured internal gasphase concentration gradients. When the base-case model parameters are used (see Table 3), the agreement between the predicted and measured radial gas concentration profiles is quite good (see Figures 6 and 7). Also, the absolute values correspond remarkably well (see Figures 5-8), considering the huge uncertainty in the kinetics used, although it has been found that, in general, the quantitative predictions of the developed model are quite sensitive to the kinetic parameters used.19 The developed reactor model also predicts the shape of the CO2 radial profile correctly, viz., nearly flat during standard gasification conditions with only primary air (see Figure 6), and important radial profiles in the case of secondary air and combustion (see Figure 8). It turns out that, for the base-case model input parameters as presented in Table 3, the measured radial concentration gradients can exist only if the radial Pe´clet number is on the order of 1000 in the dilute zone (between L ) 1.25 and 6 m), which means that the degree of radial mixing is very low.18 Also, in the lower density zone (between 0 and 1.25 m), mixing cannot be very good (say, minimal18 Per ≈ 100). In a mathematical model of a CFB reactor, it is necessary to incorporate the shape of the radial profile of the axial gas velocity, the radial solids distribution, and the rate of radial gas mixing. These phenomena are extremely important for the correct prediction of the conversion and the selectivity of chemical reactions, as can be seen from Table 4 in which results from the pilot plant are compared with model calculations for different assumptions regarding the hydrodynamics. The model results indicate that, even for biomass gasification,
which is characterized by very high rates of all dominant chemical reactions (combustion, initial tar cracking, water-gas shift), the effects of these phenomena are significant (see Table 4). It has been found that, as expected, plug flow for the gas phase and a uniform radial distribution of char and biomass improves the carbon conversion. This effect is quite important: the model predicts that the carbon conversion for an ideal situation with plug flow (of gas and solids) and a uniform radial distribution of the solids is 7% higher than for a completely segregated system [compare simulated cases 1 (base case) and 4 in Table 4]. The model presented is incomplete. It has been used here for interpretation only. On the basis of assumptions concerning radial distributions of velocities and the voidage, for which good estimates were available from experiments in this case, the observed radial gas concentration profiles can be explained. Much more information is required, however, to predict in detail the performance of CFB biomass gasifiers without any experimental input, which is, of course, the ultimate goal of modeling. 5. Improving the CFB Gasifier A multistage fluidized bed reactor has been designed and tested at the laboratory scale,20 specifically to overcome the problems usually encountered in circulating fluidized bed gasifiers. The riser of this multistage circulating fluidized bed consists of several segments (seven in the base-case design) in series, each composed of two opposite cones. The new concept creates favorable reactor conditions that cannot be realized otherwise. The benefit of a concept in which different processes are carried out in separate segments of the same reactor has been demonstrated for the specific case of biomass gasification in a laboratory setup. In this small-scale setup, it was possible to create oxidation segments in which O2 reacted exclusively with char (carbon). This resulted in an increased carbon conversion and, consequently, an improved gasification efficiency.20 Creating an exclusive char combustion zone by staging has also been tested in ECN’s pilot CFB facility. A flow restriction was introduced under the biomass feed point, which is located 1 m above the primary air nozzles. This was aimed at creating a combustion zone below the biomass feed location. The flow restriction was supposed to prevent that biomass and combustible vapors from falling back and back-mixing the combustion zone. In this way, when enough char is present, an exclusive char combustion zone is created. Gasification tests performed in the modified CFB (including the flow restriction) clearly showed increased carbon conversion and thermal efficiency compared to the conventional CFB at otherwise the same process conditions. For instance, at typical gasification equivalence ratios of ca. 0.25, the carbon conversion increased from 87 to 96%, and the thermal efficiency increased from ca. 60 to 70% on average.9,20 Scaling up these data to a commercial gasifier with almost no heat loss would lead to a thermal efficiency of 80% with nearly complete carbon conversion. 6. Conclusions CFB risers have some significant favorable characteristics, viz., a high throughput, an available heat wheel provided by the circulating solids, and a low solids
Ind. Eng. Chem. Res., Vol. 42, No. 26, 2003 6763
hold-up/pressure drop. On the other hand, the state of axial and radial gas and solids mixing is far from ideal, and the coupling between the prevailing solids circulation rate, the gas velocity (profile), and the solids holdup does not always provide well-balanced reactor conditions. For conventional CFB biomass gasifiers these adverse properties lead to the observed facts that (i) the contact between char/oxygen and tar/catalyst is not optimal, (ii) oxygen can easily react with combustible gases, and (iii) devolatilization of the fuel can take place over the whole length of the riser. To overcome these problems, a multistage circulating fluidized bed has been proposed. In this novel reactor, it is possible to combine different processes (e.g., gas-solids reactions), spatially divided, in separate segments of a single reactor and to prevent the back-mixing of gas and solids between these segments effectively. This novel reactor has been tested successfully at the laboratory scale. In addition, the concept of staging has been implemented in the riser of ECN’s conventional CFB riser resulting in a significant improvement of the thermal efficiency. Acknowledgment The authors are indebted to the ECN management for their permission to use the pilot-plant data, both the regular experiments and the special experiments carried out for this paper. Most of the data used are from sponsored tests programs. For this, the Dutch agency for energy and environment (NOVEM), the Dutch company NV Afvalzorg, and the Dutch federation of energy companies (EnergieNed) are acknowledged. The whole pilot-plant operating crew under the supervision of Christiaan van der Meijden is also acknowledged for their collaboration with this work. Notation D ) diameter of the riser, m D ) dispersion coefficient, m2/s Dsauter ) Sauter mean diameter, mm d ) diameter coordinate, m dp ) particle size (diameter), µm Gf ) gas flux, kg/(m2.s) Gs ) solids flux, kg/(m2.s) Hvap ) MJ/kg H/C ) atomic H to C ratio in the feed HHV ) higher heating value, MJ/kg or MJ/Nm3 L ) length, m LH ) latent heat, MJ/kg or MJ/Nm3 LHV ) lower heating value, MJ/kg or MJ/Nm3 p ) pressure, Pa or bar Per ) radial Pe´clet number: Per ) UR/Deff r Q ) heat loss, MJ/h or W (positive when heat is lost to the environment) R ) radius of the riser, m or cm Rgas ) gas production (including N2 and water) per kilogram of biomass, Nm3/kg r ) radial coordinate, m or cm SH ) sensible heat, MJ/kg or MJ/Nm3 T ) temperature, °C or K U ) superficial gas velocity, m/s YC(i) ) yield on a carbon basis: kg of C in component i divided by kg of C in biomass YH(i) ) yield on ahydrogen basis: kg of H in component i divided by kg of H in dry biomass Xc ) carbon conversion Xc,gas ) carbon conversion based on the carbon content in the gas phase
Xc,ashes/dust ) carbon conversion based the carbon content in the ashes and dust x ) volume fraction, % Greek Symbols R ) parameters in relation of Godfroy et al.14 ) voidage ηLHV ) thermal efficiency γ ) parameter in the relation of Godfroy et al.14 λ ) equivalence ratio Θ ) flow rate, kg/s or Nm3/s θ ) dimensionless radial coordinate (r/R) F ) density, kg/m3 σ ) standard deviation Subscripts and Superscripts bio ) biomass c ) carbon eff ) effective m ) mass p ) particle th ) thermal r ) radial coordinate v ) volume z ) axial coordinate Operators 〈j〉 ) cross-sectional average value of j 〈〈j〉〉 ) volumetric average value of j Abbreviations BFB ) bubbling fluidized bed CFB ) circulating fluidized bed daf ) dry ash free vol ) volume wt ) weight
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Received for review October 18, 2002 Revised manuscript received October 3, 2003 Accepted October 3, 2003 IE020818O
