Air–steam gasification of biomass in a fluidised bed: Process optimisation by enriched air

Fuel Processing Technology 90 (2009) 677–685

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Fuel Processing Technology j o u r n a l h o m e p a g e : w w w. e l s ev i e r. c o m / l o c a t e / f u p r o c

Air–steam gasification of biomass in a fluidised bed: Process optimisation by enriched air Manuel Campoy, Alberto Gómez-Barea ⁎, Fernando B. Vidal, Pedro Ollero Bioenergy Group, Chemical and Environmental Engineering Department, Escuela Superior de Ingenieros (University of Seville), Camino de los Descubrimientos s/n. 41092 - Seville, Spain

a r t i c l e

i n f o

Article history: Received 8 August 2008 Received in revised form 4 December 2008 Accepted 12 December 2008 Keywords: Gasification Biomass Fluidised-bed Scaling-up Enriched air–steam mixtures

a b s t r a c t The effect of oxygen concentration in the gasification agent was studied by enriched–air–steam biomass gasification tests in a bubbling fluidised-bed gasification (FBG) plant. The oxygen content in the enriched air was varied from 21% (v/v, i.e. air) to 40% (v/v), aiming at simulating FBG where enriched air is produced by membranes. The stoichiometric ratio (ratio of actual to stoichiometric oxygen flow rates) and steam-to-biomass ratio (ratio of steam to biomass, dry and ash-free, flow rates) were varied from 0.24 to 0.38 and from 0 to 0.63, respectively. The tests were conducted under simulated adiabatic and autothermal conditions, to reproduce the behaviour of larger industrial FBG. The temperature of the inlet gasification mixture was fixed consistently at 400 °C for all tests, a value that can be achieved by energy recovery from the off-gas in large FBG without tar condensation. It was shown that the enrichment of air from 21 to 40% v/v made it possible to increase the gasification efficiency from 54% to 68% and the lower heating value of the gas from 5 to 9.3 MJ/Nm3, while reaching a maximum carbon conversion of 97%. The best conditions were found at intermediate values of steam-to-biomass ratio, specifically within the range 0.25–0.35. The enriched-air–steam gasification concept explored in this work seems to be an interesting option for the improvement of standalone direct air–blown FBG because it considerably improves the process efficiency while maintaining the costs relatively low as compared to oxygen-steam gasification. © 2008 Elsevier B.V. All rights reserved.

1. Introduction Gasification is a promising technology for biomass and waste utilisation with low environmental impact, reducing global CO2 emissions. Standalone air-blown bubbling fluidised-bed gasification (FBG) is the simplest and probably the most cost-effective concept for medium-scale thermal and electricity applications (b5–10 MWe). It has been successfully demonstrated connected to a large coal-fired boiler in power plants, resulting in high efficiency when burning biomass. For small-scale power production a unit comprised of a gasifier and a compression-ignition engine is less expensive than a boiler-based power cycle, thus providing an attractive option for remote locations [1]. The use of air in gasification yields a fuel gas highly diluted by nitrogen. The lower heating value (LHV) of the gas is therefore typically below 6 MJ/Nm3. The use of steam has proven effective for achieving a medium heating value (up to 14 MJ/Nm3), but the process becomes more complex. Two different concepts have been developed, for full-scale steam gasification: steam-oxygen mixtures and indirect gasification based on twin-bed reactors [2].

⁎ Corresponding author. Tel.: +34 95 4487223; fax: +34 95 4461775. E-mail address: [email protected] (A. Gómez-Barea). 0378-3820/$ – see front matter © 2008 Elsevier B.V. All rights reserved. doi:10.1016/j.fuproc.2008.12.007

In steam-oxygen gasification the biomass is partially oxidised to provide the heat necessary to make the process self-sufficient. The gas produced has a high hydrogen content and the dilution of nitrogen is avoided. This concept has been extensively tested at laboratory scale [3,4], but the high cost of pure oxygen (for instance, based on distillation units) makes the implementation of this process uncertain at industrial scale [5]. Indirect biomass gasification is based on the separation of the gasification and combustion stages in two different chambers. The heat supply to the gasification process, which is fed with steam, is generated in the combustion chamber, in which the char is burnt out using air. The heat transfer is achieved by circulation of the bed inventory between the two stages. The essential contribution of this concept is that it allows ‘autothermal’ steam gasification without the need of oxygen (only air is used), producing a gas with medium heating value, i.e., not diluted by nitrogen. This technology has achieved semi-commercial status today [6,7]. Much experimental work has been done on FB biomass gasification using different gasification agents. Air [8], pure steam [9–11], oxygensteam [3,4] and air–steam [12–14] tests have been carried out in various lab-scale facilities, generating useful knowledge for understanding the process. Relatively little work has been found, however, using enriched–air–steam mixtures as gasification agent [15]. Moreover, most studies at laboratory or pilot scale have been conducted allothermally; the temperature, air/oxygen-to-biomass ratio and

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steam-to-biomass ratio have been varied independently because the temperature of the gasifier is controlled by external heat addition using electric heaters. However, this method of supplying heat is neither technically nor economically feasible for large-scale implementation, so the results from “allothermal” lab-scale rigs must be interpreted with caution. In previous work [16], we have studied the process improvement by using air–steam mixtures under simulated adiabatic and autothermal conditions. We have shown how the optimisation of air– steam mixtures in autothermal FBG increases the gasification efficiency from 40% to 60%, while maintaining the low heating value of the gas at around 5 MJ/Nm3. The process was, however, limited to using air as the source of oxygen to the system, i.e., 21% (v/v). In this study we have extended the previous work by allowing the oxygen to vary in purity up to 40% (v/v) in the enriched air. This higher figure was chosen because it is the oxygen concentration that can be produced using commercial air separators based on membrane technologies. This way of producing oxygen would keep the investment and operational costs relatively low compared to processes that need pure oxygen, usually based on distillation units. Consequently, the size of plants where this concept is expected to be feasible is not necessarily large. Therefore, the concept investigated in this work could be convenient for medium-scale electricity production (b5–10 MWe), provided the gas cleaning, particularly tars, is done properly. This study looked at gas quality and the impact on carbon conversion and process efficiency. Tar production and cleaning strategies will be addressed in future work. The tests were conducted varying the flow rates of biomass, steam and enriched air and the oxygen concentration in the enriched air, making it possible to explore the effects of these variables on the quality and composition of the produced gas, aiming at process optimisation. All the tests were conducted with wood pellets in a pilot-scale FBG simulating autothermal and adiabatic conditions. In addition, the temperature of the gasification agent was fixed at 400 °C for all tests, a value that can be achieved by energy recovery from the off-gas without tar condensation. With this experimental setup, we expect that the results will be useful for scaling up the data to design industrial FBG, i.e., large standalone direct gasification units without external heat supply and with low wall-heat loss. 2. Experimental 2.1. Materials The biomass used was wood pellets with the empirical formula CH1.4O0.64 (dry and ash-free, d.a.f.) [16]. The moisture and ash content were 6.3% and 0.5% (w/w), and the lower heating value of the fuel (as received) was 17.1 MJ/kg. The pellets were cylindrically shaped with a mean diameter of 6 mm and 5–10 mm long. The apparent density of the pellets and the bulk density were 1300 and 600 kg/m3. The bed material used was ofite, a sub-volcanic rock composed mainly of feldspar, pyroxene and limestone, whose complete characterisation is given in [17]. 2.2. Facility Fig. 1 shows the layout of the facility and Table 1 gives the main parameters of the pilot plant. The rig has been described in detail in a previous publication [16]. The only modifications made in the plant with respect to previous work were the erection of a plant for oxygen production and the removal of the tar scrubber to avoid tar condensation in the pipes between the cyclones and the combustion chamber. In addition, the pipes were maintained at a temperature above 400 °C by heating elements and insulation blankets. The oxygen plant comprised four 10-m3 oxygen bottles, making it possible to produce 10 Nm3/h with a purity higher than 99%.

2.3. Test procedure The protocol for the tests was similar to that used in [16]. Therefore, only a brief description is given here. At the beginning of each test a batch of bed material (around 8 kg) was added to the reactor. The bed was heated up with the hot air and the electric heater. The bed and the freeboard were rapidly heated up to approximately 700 °C. It was necessary, however, to wait longer before starting the biomass feed in order to avoid tar deposition on the pipes between the cyclones and the combustion chamber. Once the temperature upstream of the combustion chamber was higher than 300 °C, the facility was considered ready for biomass, and biomass was fed slowly into the reactor. In these conditions of oxygen excess, the biomass was oxidised completely and the reactor was rapidly heated to the desired process temperature. From the beginning of the test, a computer-based data acquisition system monitored and recorded the temperatures, pressures, gas composition (H2, CO, CO2, CH4, O2), power supplied to the heating equipment and the flow rates of gas and solid. The transition from combustion to gasification was made by increasing the biomass flow rate to decrease the air-to-biomass ratio. The steam addition started once a steady-state condition was established. The mixture of steam and air was heated in a preheater and then fed to the FBG. Once the process condition was stable, the oxygen was added to the steam before mixing with air, with the flow rate reduced to an appropriate value to obtain the desired level of oxygen in the final mixture. The final air–steam–oxygen mixture was heated in the preheater to the process temperature (400 °C). In all tests an initial transitory period of 3 to 4 h was followed by a steadystate period of 5 to 7 h. The operation was finalised by taking a sample from the bed inventory and combusting the remaining char in the bed. After each test, the two cyclone bins and the extraction-ash bin were sampled and analysed. 2.4. Operating conditions The mode of operation we used for conducting the tests (adiabatic tests and temperature of inlet gasification mixture fixed at 400 °C) for a given biomass (wood pellets) consisted of four variables that could be independently varied: the flow rates of biomass, air, steam and oxygen. The flow rates of air and oxygen determined the oxygen content in the enriched air inlet, i.e., the O2 purity (hereafter referred to as OP, in % v/v). For a given biomass flow rate, two ratios can be defined for the analysis of the process: (1) the stoichiometric ratio (SR), defined as the mass ratio between the amount of total oxygen fed in and the stoichiometric amount of oxygen required for combustion, and (2) the steam-to-biomass ratio (SBR), defined as the flow rate of steam fed to the reactor divided by the biomass flow rate (dry and ash-free). Two additional variables must be taken into account: (3) the oxygen percentage of the enriched air (OP), which is an indication of the nitrogen dilution of the produced gas, and (4) the biomass throughput. However, for a limited range of this parameter in an FB, the system can be analysed approximately by a mean of (1), (2) and (3), as the biomass flow rate is expected to have a minor influence on the results. The experimental program comprised tests with air–oxygen– steam mixtures in different proportions. In the first set of tests, air was used as oxygen supplier, so the OP was 21% v/v for all tests and the oxygen flow rate was nil. Some of these tests have been reported in previous work [16]. The second set of tests involved air–oxygen-steam mixtures in which the OP of the enriched air was set at 30, 35 and 40% v/v. Some of these tests were conducted twice to ensure the reproducibility of the results. The experimental programme comprised tests varying SR, SBR and OP between the indicated ranges: SR from 0.24 to 0.38, SBR from 0 to 0.63 and OP from 21 to 40%. To establish these ratios, the following flow rates were varied between the indicated ranges: the steam flow

M. Campoy et al. / Fuel Processing Technology 90 (2009) 677–685

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Fig. 1. Pilot plant layout.

rate from 0 to 7.4 kg/h, the biomass flow rate from 10 to 21.6 kg/h, the air flow rate from 5.6 to 17 Nm3/h and the oxygen flow rate from 0 to 3.7 Nm3/h. An experimental matrix was constructed with four values of OP (21, 30, 35 and 40) and two levels of SR and SBR, representing low and high values: 0.24–0.27 (low SR) and 0.33−0.38 (high SR) for SR and 0.22–0.36 (low SBR) and 0.43–0.63 (high SBR) for SBR. The experimental conditions were selected to assess the operation at low and high values of both SR and SBR, while changing the OP to keep the temperature high enough. Table 2 groups the tests in four combinations on the basis of the level of the SR and SBR used. In total, the matrix of experiments comprised 16 independent experiments. The main results of the tests are given in Table 3. Table 3 includes 20 tests because four additional tests were conducted to study the effect of steam in more detail. These additional tests were conducted at a fixed value of SR ≈ 0.35 and two OP values (21% in test 1 and 40% in tests 14, 16 and 17). The gas residence times in the bed and the freeboard are important operating data, and should therefore be specified clearly. Ideally, the gas residence time in the bed and the freeboard should take into account the gas from biomass devolatilisation as well as that from the gasification of char. In this work, we calculated the residence time in

Table 1 Technical and operating data for the facility and the main test operating data. Inside Bed Diameter Inside Freeboard Diameter Freeboard Height Bed Material Fuel Fuel Flow Rate Gasification Agent Operation Temperature Fluidisation Regime

0.15 m 0.25 m 2.15 m Ofite Wood Pellets 10–21.6 kg/h Air + Oxygen + Steam 755–840 °C Bubbling

the bed based on the total amount of gasification agent fed to the reactor at the mean bed temperature. For the gas residence time in the freeboard, the calculation was made by considering the outlet dry gas and the average freeboard temperature (water content was not considered because of the uncertainty of this value).The gas residence times for the bed and the freeboard were calculated taking into account the dimensions given in Table 1 and ranged from 0.99 to 1.43 s and between 3.7 and 6.1 s, respectively. One significant aspect of the tests described in this work is that they were conducted nearly adiabatically: the heating system was controlled to provide just the necessary amount of heat to compensate for the heat losses. This operation was achieved by keeping the temperature of the furnace slightly lower (5–10 °C) than the temperature measured inside the reactor. This operational procedure has been explained in [16]. In addition, the temperature of the gasification agent was fixed at 400 °C for all tests, a value which can be achieved by heat recovery in standalone FB systems without tar condensation. 3. Results and discussion The variables analysed included gas composition, gas yield, heating value, gasification efficiency and carbon conversion.

Table 2 Definition of the four combinations used for the analysis as a function of the levels of SR and SBR tested. SR

SBR

Combination

High

Low

High

Low

1 2 3 4

0.33–0.38 0.33–0.38 – –

– – 0.25–0.27 0.24–0.27

– 0.45–0.60 – 0.43–0.63

0.22–0.36 – 0.23–0.31 –

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Table 3 Test results. 1

2

3

4

5

6

7

8

9

10

11

12

13

14

15

16

17

18

19

20

Operational conditions Biomass Flow Rate (kg/h) Air Flow Rate (Nm3/h) Oxygen Flow Rate (Nm3/h) Steam Flow Rate (kg/h) Mean Bed Temperature, Tb (°C) Furnace Set-point (°C) Mean Freeboard Temperature(°C) tbed (s)⁎ tfreeboard (s)⁎

11.5 17 0 0 812 800 716 1.3 4.3

12.2 17 0 2.5 804 800 721 1.1 4.0

12.2 17 0 5.1 789 780 709 1.0 4.0

15 17 0 3.2 786 780 708 1.1 3.9

15 17 0 6 755 750 709 1.0 3.7

12.4 11.9 1.5 3.7 808 800 715 1.2 4.7

10.0 9.1 1.2 5.6 790 780 715 1.3 6.1

16.2 10.6 1.4 4.7 781 770 716 1.3 4.5

12.0 7.7 1.0 6.5 765 750 695 1.4 6.0

14 11.6 2.3 4.3 820 810 715 1.2 4.2

11.8 8.3 1.8 6.2 795 780 725 1.3 5.7

16.8 9.5 2 4.9 800 790 725 1.3 4.5

12.6 7 1.5 7.4 757 750 720 1.3 5.8

21.6 11.5 3.7 2.1 840 825 724 1.2 3.4

16.2 9.3 2.9 4.4 829 815 708 1.2 4.4

14.8 8.9 2.8 4.9 830 815 727 1.2 4.7

13.2 7.3 2.3 6.6 813 800 725 1.3 5.5

12.0 6.8 2.1 6.4 806 800 727 1.3 5.8

18.8 8.1 2.6 5.3 803 790 725 1.3 4.4

14.0 5.6 1.8 7.3 766 760 722 1.4 6.0

Main variables for analysis OP (%) SR SBR Combination (see Table 2)

–

Gas: composition and yields CO (%v/v, dry) H2 (%v/v, dry) CO2 (%v/v, dry) CH4 (%v/v, dry) CO (%v/v, dry, N2 free) H2 (%v/v, dry, N2 free) CO2 (%v/v, dry, N2 free) CH4 (%v/v, dry, N2 free) CO yield (g/kg d.a.f.b.) H2 yield (g/kg d.a.f.b.) CO2 yield (g/kg/d.a.f.b.) CH4 yield (g/kg d.a.f.b.)

15.8 8.7 15.1 5.1 35.3 19.5 33.8 11.4 453.4 17.8 680.9 83.6

Process variables Gas Flow Rate (Nm3/h, dry)⁎ Gas yield (GY) (Nm3 dry gas, N2 free/kg d.a.f.b.) Low healing value (LHV) (MJ/Nm3 dry gas) Cold Gasification Efficiency Carbon Conversion

21 0.35 0

24.6 1.03 4.76 0.59 0.93

⁎ Not measured (calculated from other measurements).

21 0.33 0.22 1

15.4 11.9 15.9 4.8 32.1 24.8 33.1 10.0 443.4 24.5 719.3 79.0

26.1 1.11 4.95 0.62 0.90

21 0.33 0.45 2

13.8 13.3 17.0 4.6 28.3 27.3 34.9 9.4 402.7 27.7 779.6 76.7

26.5 1.14 4.83 0.61 0.92

21 0.27 0.23 3

15.0 14.0 16.2 4.7 30.1 28.1 32.5 9.4 365.7 24.4 620.6 65.5

27.2 0.97 5.09 0.54 0.90

21 0.27 0.43 4

11.9 16.2 18.6 5.3 22.9 31.2 35.8 10.2 302.8 29.4 743.7 77.1

28.4 1.06 5.15 0.57 0.91

30 0.36 0.32 1

18.9 16.4 17.6 5.5 32.3 28.1 30.2 9.4 456.1 28.3 669.7 76.0

22.3 1.13 6.12 0.64 0.94

30 0.35 0.60 2

15.7 18.3 18.8 5.7 26.9 31.2 32.2 9.7 362.8 30.1 681.6 75.1

17.2 1.08 6.00 0.60 0.95

30 0.25 0.31 3

20.8 20.0 15.8 6.7 32.9 31.6 24.9 10.6 405.0 27.8 482.0 74.5

23.5 0.98 7.19 0.61 0.96

30 0.24 0.58 4

15.3 22.3 20.3 7.1 23.5 34.3 31.3 10.9 305.6 31.9 639.3 81.2

17.9 1.04 6.88 0.60 0.96

35 0.38 0.33 1

20.0 17.5 16.8 5.6 33.4 29.2 28.0 9.4 475.1 29.7 626.7 76.1

24.8 1.14 6.41 0.66 0.95

35 0.34 0.56 2

17.5 21.8 18.0 6.1 27.6 34.4 28.4 9.6 366.1 32.6 591.5 72.9

18.4 1.06 6.75 0.61 0.95

35 0.27 0.31 3

23.9 22.4 12.6 7.3 36.1 33.9 19.0 11.0 438.1 29.4 361.8 76.5

22.9 0.97 8.06 0.64 0.96

35 0.26 0.63 4

19.3 25.1 16.2 7.4 28.4 36.9 23.8 10.9 369.3 34.3 487.0 80.7

17.9 1.04 7.81 0.65 0.96

40 0.32 0.10 –

27.4 18.3 16.2 7.3 39.6 26.5 23.4 10.5 520.1 24.9 481.9 79.2

30.5 1.05 8.06 0.67 0.97

40 0.33 0.29 1

25.1 23.1 13.7 6.5 36.7 33.8 20.0 9.5 499.4 32.9 427.6 73.9

24.0 1.09 8.00 0.69 0.96

40 0.35 0.36

40 0.32 0.54

–

–

23.9 22.3 14.6 6.7 35.4 33.0 21.6 9.9 483.5 32.2 464.4 77.5

20.2 24.5 16.7 6.9 29.6 35.9 24.4 10.1 386.7 33.5 501.5 75.4

22.3 1.09 7.83 0.69 0.96

18.8 1.05 7.67 0.64 0.97

40 0.33 0.57 2

19.3 25.7 17.0 6.7 28.1 37.4 24.8 9.7 384.7 36.6 532.7 76.2

17.8 1.10 7.62 0.66 0.96

40 0.26 0.30 3

28.5 25.7 9.2 8.1 39.8 36.0 12.9 11.3 476.2 30.8 242.0 77.5

23.4 0.96 9.28 0.68 0.97

40 0.24 0.56 4

23.5 27.5 14.6 7.7 32.1 37.5 19.9 10.5 391.5 32.7 381.3 73.2

17.3 0.98 8.70 0.63 0.97

M. Campoy et al. / Fuel Processing Technology 90 (2009) 677–685

Test

M. Campoy et al. / Fuel Processing Technology 90 (2009) 677–685

Fig. 2. (a) Bed temperature as a function of OP for different SR and SBR levels. : SR= High, SBR = Low (combination 1); : SR= High, SBR = High (combination 2); : SR= Low, SBR = Low (combination 3); SR= Low, SBR = High (combination 4). (See Table 2 for the definition of the four combinations). (b) Comparison between mean bed temperature (Tb, ) and the theoretical bed temperature (T⁎, ). (c) Bed temperature as b solid line and a function of SBR for SR= 0.35 and different levels of OP: : OP= 21%; : OP= 30%; : OP= 35%; OP= 40%.

The effect of OP on the gasification temperature is shown in Fig. 2(a), indicating that the temperature in the bed increased with OP in the enriched air for all SR and SBR levels analysed. For fixed values of SR and SBR, the increase in OP means that the same amount of combustible matter is burned whilst the nitrogen flow that had to be heated decreased, leading to a higher bed temperature. However, our measurements did not identify such as increase in temperature that would be expected from merely a decrease in the nitrogen dilution. We demonstrate this observation by means of Fig. 2(b). This figure shows the measured (average) bed temperature, Tb, for four tests (solid diamonds) and Tb⁎ (blank squares over the solid line) as a function of OP, where Tb⁎ is the bed temperature for those tests considering that the only effect of decreasing the nitrogen flow rate was to raise the bed temperature. Specifically, Fig. 2(b) presents the values of Tb and Tb⁎ for combination 1 (high SR and low SBR). For the calculation of Tb⁎ for a test with OPN 21%, say test 6, a heat balance was made by assuming that

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the difference between the heat for heating up the fluidisation agent in that test (6) as compared to that in test 2 (reference), was entirely in the form of sensitive energy of the outlet gas. The same assumption was made for the calculation of the corresponding Tb⁎ for tests 10 and 15 in Fig. 2(b). Then, the fact that, for a fixed OP, in Fig. 2(b) Tb is lower than Tb⁎ is interpreted as an enhancement of endothermic processes such as tar and hydrocarbon reforming as well as char gasification with CO2 and H2O. These processes are kinetically-limited under the operating conditions tested, so a rise in the temperature enhances the rate of these reactions. The actual bed temperature in a test with OP N 21% is the net result of two opposing processes, one promoting an increase in the bed temperature while the other making the temperature lower. The difference between Tb and Tb⁎ for the tests with OP N 21% in Fig. 2(b) confirmed that the aforementioned endothermic processes are significant. This is confirmed below by observing an increase of the CO and H2 concentration, expressed in nitrogen-free basis, with OP. Fig. 2(c) shows the effect of SBR on gasification temperature using various degrees of OP for a fixed value of SR ≈ 0.35. As expected, the temperature in the bed decreased as the steam flow rate increased. This occurred for the four values of OP studied. The effect of OP in the enriched air on the gas composition is given in Fig. 3 (% v/v of dry gas) and Fig. 4 (yields per kilogram of dry and ash-free biomass, d.a.f.b.). The volume fraction of the main gas species measured in the outlet gas is shown in Table 3, expressed in dry basis and also in dry nitrogen-free basis. As seen the composition of CO, H2, CO2 and CH4 presents substantial differences depending on the basis used. For a given combination, the molar composition of CO and H2 increased with OP, expressed in both dry and dry and nitrogen-free basis. This fact supports the conclusion made above that the nitrogen dilution is not the only effect caused by the increasing of OP. Fig. 3(a) shows that, for all values of SR and SBR, the carbon monoxide increased as OP increased. An increase of carbon monoxide from 12 up to 28 was observed when OP was changed from 21 to 40%. The combination of low SR and low SBR (combination 3 defined in Table 2), seemed to provide the largest amount of CO, whilst combination 2 (high SR and SBR) led to lower CO concentration in the gas. The CO content increased as the OP rose. The H2 content in the gas is shown in Fig. 3(b). It is clearly seen that combination 4 (low SR and high SBR; see Table 2) increased the hydrogen content in the gas for all values of OP tested. Fig. 4 shows the yields of carbon monoxide and hydrogen per kilogram of d.a.f.b. of CO and H2, respectively, as a function of the OP for the different combinations of SR and SBR. We note that both yields increased with OP. The yield of methane was roughly constant (75 g/kg d.a.f.b.) These findings are quite reasonable because, in general, as SR is increased, a smaller amount of CO ends up in the produced gas and more nitrogen is added to the system. In low-temperature gasification the devolatilisation of the biomass determines the resulting gas composition to a large extent. CO is oxidised to a larger extent, resulting in lower CO yields. At a fixed SR, increasing OP means less nitrogen in the gas and hence higher CO concentration for the same CO yields. In addition, the rising temperature can increase CO yields for a given SR and SBR (as seen in Fig. 4(a)), in this way making the effect of increasing CO concentration even greater. The fact that SBR increased the H2 content is also reasonable, but the behaviour with respect to SR was not known a priori; less hydrogen was probably burned as the SR decreased. However, for a constant OP and SBR, the thermal level in the gasifier was lower, resulting in a lower gas yield of H2 and possibly less secondary gas phase reaction, supporting the presence of H2, such as in cracking and reforming of hydrocarbons and tars. Secondary gas reactions such as the water-gas-shift-reaction (WGSR) can alter the final composition if

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and oxidation of carbon monoxide. In fact, the situation is more complex since the CO2 is also a reactant for heterogeneous reactions with char and possibly with tar; hence there is competition for the oxygen among different species present in the gasifier: hydrogen, char, tar, methane, light hydrocarbons, etc. It is most likely that hydrogen and carbon monoxide react with oxygen to a greater extent since they are more reactive species. This conjecture is also supported by the small variation detected for methane in Fig. 3(d). A methane increase of only up to 3% was observed as OP increased from 21 to 40%. The complex processes occurring simultaneously can explain the trend shown in Fig. 3(c) for low values of OP, typically representing conditions at lower temperature. As seen, the CO2 concentration was slightly affected by the OP at low values of this parameter. From an overall comparison of different graphs in Fig. 3, it seems that devolatilisation and subsequent hydrogen and carbon monoxide oxidation provide a good explanation for the gas distribution measured at the exit of the gasifier. The four graphs in Fig. 3 establish that CO and H2 increased, CO2 decreased and CH4 decreased slightly with the OP in the enriched air. The WGSR does not explain the gas phase distribution observed since it seems that the hydrogen and carbon monoxide increased for all values of OP. Other homogeneous, and to a lesser extent heterogeneous, reactions probably determined the gas product distribution and the equilibrium of the WGSR was not attained. In general, the reduction of the nitrogen dilution effect that took place when increasing the flow rate of oxygen also explains the trends observed. The gas yield (GY), defined as the ratio of the volumetric flow rate of the dry and nitrogen-free product gas (without considering light hydrocarbons and tars) and the mass flow rate of biomass (d.a.f.), is shown in Fig. 5. As observed, there was not a clear increase in gas yield with OP for any of the SR-SBR tested. It would be expected, however, that the gas yield would increase with OP, since for the same SR and SBR, the temperature attained in the gasifier was higher and,

Fig. 3. Gas composition as a function of OP for different SR and SBR levels. : SR=High, SBR=Low (combination 1); : SR=High, SBR=High (combination 2); : SR=Low, SBR=Low (combination 3); SR=Low, SBR=High (combination 4). (a) carbon monoxide, (b) hydrogen, (c) carbon dioxide and (d) methane.

the temperature and residence time of the gas are sufficiently high. This is analysed below by comparing the simultaneous effects on all the species. Fig. 3(c) shows the effect of OP on the carbon dioxide, evidencing behaviour compatible with the behaviour of carbon monoxide shown above; carbon dioxide is the result of both biomass devolatilisation

Fig. 4. Gas yield as a function of OP for different SR and SBR levels. : SR=High, SBR=Low (combination 1); : SR=High, SBR=High (combination 2); : SR=Low, SBR=Low (combination 3); SR=Low, SBR=High (combination 4). (a) carbon monoxide and (b) hydrogen.

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Fig. 5. Gas yield as a function of OP for different SR and SBR levels. : SR = High, SBR = Low (combination 1); : SR = High, SBR = High (combination 2); : SR = Low, SBR = Low (combination 3); SR = Low, SBR = High (combination 4).

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of oxygen) is intended to be used for burning in thermal applications or, if the gas is properly cleaned, for power production. The LHV of the gas must consistently include the gross gas, including nitrogen. This is not the usual way to calculate the LHV in most of air–steam gasification works, where the LHV is expressed on a nitrogen-free basis because the gas is mainly aimed at medium-heating-value gas production or for high-hydrogen gas production. Gasification efficiency, defined as the ratio of combustion heat of the produced gas calculated using the LHV defined above, to combustion heat of the biomass, is shown in Fig. 7. Cold gasification efficiency assumes a temperature of the product gases at 25 °C, so the sensitive heat of the gas is not taken into account. Fig. 7(a) displays the cold gasification efficiency as a function of OP for the two levels of SR and SBR investigated, showing that the increase of OP in the enriched air increased the efficiency from 55% up to 70%, the latter corresponding to the higher SR tested. Fig. 7(b) depicts the effect of

consequently, more gas was produced via more intense devolatilisation. In addition, enhanced tar decomposition and char gasification would be expected, which would contribute to further increasing the gas yield. The behaviour observed in Fig. 5, however, probably has to do with the way that the gas yield was calculated. It included the contribution of the four main species measured online: CO, CO2, H2 and CH4. Fig. 6 shows the lower heating value of the product gas (LHV), defined as the heat of combustion at 25 °C of 1 Nm3 of gas, as a function of the OP for the four combinations of SBR and SR tested. As expected, the higher the OP, the higher the LHV of the gas. The analysis of gas composition in Figs. 3 and 4 explains the behaviour observed for the LHV in Fig. 6. Low SR leads to more combustible gas species in the gas produced, whereas higher OP makes it possible to reach a higher temperature in the gasifier with the same degree of fuel oxidation, i.e., SR level. Low SR and high OP therefore lead to enhancement of reforming reactions. In addition, it is expected that, up to some values of SBR, not known a priori, the higher concentration of steam leads to higher LHV, because the presence of steam enhances the reforming reactions even more if a sufficiently high temperature is maintained in the reactor. We can clearly envisage an optimised operation window by proper synchronisation of the various input parameters. The following are a few comments on the numerical values of LHV given in this work. On the one hand, the LHV is calculated taking into account the energy content of H2, CO and CH4. Other fuel gases such as C2 and other light hydrocarbons (which were not measured), as well as tar, were not considered in the computed LHV, so the actual LHV of the gas should be higher than reported here. On the other hand, the LHV of the dry gas, includes the presence of nitrogen, because the gas produced in this process (an autothermal gasifier using air as a source

Fig. 6. LHV of the product gas as a function of OP for different SR and SBR levels. : SR=High, SBR=Low (combination 1); : SR=High, SBR=High (combination 2); : SR=Low, SBR=Low (combination 3); SR=Low, SBR=High (combination 4).

Fig. 7. Cold gasification efficiency under different process conditions. (a) Cold gasification efficiency as a function OP for different SR and SBR levels. : SR=High, SBR=Low (combination 1); : SR=High, SBR=High (combination 2); : SR=Low, SBR=Low (combination 3); SR=Low, SBR=High (combination 4); (b) Cold gasification efficiency as a function of OP for SR≈0.35 at three levels of SBR: : 0-0.2, : 0.2-0.4 and : 0.4-0.6; (c) Cold gasification efficiency as a function of SBR for SR≈0.35 at four OP: : OP=21%; : OP=30%; : OP=35%; OP=40%.

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in carbon conversion in Fig. 8 was higher in the lower range of OP, from 21% to 30%. Fig. 8(b) shows the effect of SBR on carbon conversion for the four OP studied, demonstrating that the use of steam affects the carbon conversion only slightly. This result agrees with the analysis of the efficiency optimum given in Fig. 7. 4. Conclusions

Fig. 8. Carbon conversion under different process conditions. (a) Carbon conversion as a function of OP for different SR and SBR levels. : SR= High, SBR = Low (combination 1); : SR= High, SBR = High (combination 2); : SR= Low, SBR = Low (combination 3); SR= Low, SBR = High (combination 4); (b) Carbon conversion as a function of SBR for SR≈ 0.35 at four OP: : OP= 21%; : OP= 30%; : OP= 35%; OP= 40%.

The effect of oxygen concentration in the gasification agent was studied experimentally in a bubbling fluidised-bed wood gasifier (FBG). In this study we have extended previous work on steam-air gasification in which the oxygen content was supplied by air (21% (v/v)), by allowing the oxygen in this case to vary in purity up to 40% (v/v) in the enriched air. The choice of greater oxygen purity (40% (v/v)) was made because that is the oxygen concentration that can be produced using commercial air separators based on membrane technologies. Two different levels of SR and SBR were tested (high and low), in order to evaluate the effect of enriched air for different process conditions. The use of enriched air reduces the nitrogen dilution effect, increasing the gasification temperature. This allows the addition of low-quality steam while maintaining the thermal level in the gasifier. The appropriate combination of temperature and steam leads to higher CO and H2 yields, heating value, carbon conversion and gasification efficiency. The optimisation of steam addition using enriched air leads to a maximum efficiency of 70% for steam-tobiomass ratio around 0.3 at 40% oxygen purity. The gasification concept studied in this work seems to be an interesting option for standalone direct air–blown FBG, considerably improving the process efficiency while keeping the costs relatively low as compared to oxygen-steam gasification. Acknowledgments

the OP over cold gasification efficiency for a given value of SR ≈ 0.35 at three SBR levels: low (0–0.2), medium (0.2–0.4) and high (0.4–0.6). The increase in the cold gasification efficiency for all levels of SBR tested is clearly shown. The average value of efficiency goes from 60% (for air, i.e., OP = 21%) to 67% (OP = 40%). Cold gasification efficiency as a function of SBR with different OP is given in Fig. 7(c), which shows that the use of steam increased the efficiency up to 70% for OP = 40%. Although there are not enough tests to draw more precise trend curves, it is observed that the efficiency reached a maximum. This was argued previously when analysing the effects of process conditions on gas composition and yields. In particular, a maximum is observed in Fig. 7 (c) for OP = 40%. The cases of OP = 21, 30 and 35% suggest similar behaviour, though the lack of points does not allow us to define clear trends. Despite this uncertainty, Fig. 7(c) suggests that there is a threshold of SBR at which the presence of steam seems to lower the efficiency. The appearance of maximum efficiency at intermediate values of SBR is explained by the compromise between the enhancement of the gasification reactions and the reduction of the bed temperature that occurs as steam addition is increased. This result suggests that there is an optimal steam flow rate, which seems to be in the range 0.25–0.35 for SR = 0.35 and OP of 40%. Fig. 8(a) illustrates carbon conversion, calculated as the difference between the carbon flow rate in the feed and cyclone ash, divided by the flow rate of carbon in the feed. The flow rate of carbon for both streams is computed by the product of the stream (fuel and ash) flow rate and the total carbon content determined by ASTM analysis. It is shown that carbon conversion increases with OP for all SBR and SR levels, ranging from 91 to 97%. The behaviour shown in Fig. 8 is explained by the improvement of the gasification reactions at high temperature, leading to greater char conversion. This is supported by the analysis of the ash collected in the cyclones. The increase

This work was partly financed by the European Commission, the Commission of Science and Technology of Spain and the Regional Government of Andalusia (Junta de Andalucía). The authors also acknowledge Antonio Albea and Antonio Cabello for their assistance in performing the experimental work. References [1] E. Kurkela, M. Nieminen, P. Simell, Development and Commercialisation of Biomass and Waste Gasification Technologies from Reliable and Robust Co-firing Plants Towards Synthesis Gas Production and Advanced Power Cycles, Proceedings of the 2nd World Conference on Biomass for Energy, Industry and Climate Protection, Rome, Italy, 10-14 May 2004. [2] J. Corella, J.M. Toledo, G. Molina, A review on dual fluidized-bed biomass gasifiers, Ind. Eng. Chem. Res. 46 (2007) 6831–6839. [3] J. Gil, M.P. Aznar, M.A. Caballero, E. Francés, J. Corella, Biomass gasification in fluidized bed at pilot scale with steam-oxygen mixtures. product distribution for very different operating conditions, Energy Fuels 11 (1997) 1109–1118. [4] Y. Wang, C.M. Kinoshita, Experimental analysis of biomass gasification with steam and oxygen, Sol. Energy 49 (1992) 153–158. [5] K. Maniatis, Progress in biomass gasification: an overview, in: A.V. Bridgwater (Ed.), Progress in Thermochemical Biomass Conversion, Blackwell Science, London, 2001, pp. 1–31. [6] M.A. Paisley, R.P. Overend, The SylvaGas Process from Future Energy Resources - A Commercialization Success, 12th European Biomass Conference, Amsterdam, The Netherlands, 17-21 June 2002. [7] R. Rauch, Steam Gasification of Biomass at CHP Plant in Güssing - Status of the Demonstration Plant, Proceedings of the 2nd World Conference on Biomass for Energy, Industry and Climate Protection, Rome, Italy, 10-14 May 2004. [8] I. Narváez, A. Orio, J. Corella, M.P. Aznar, Biomass gasification with air in an atmospheric bubbling fluidized bed. effect of six operational variables on the quality of the produced raw gas, Ind. Eng. Chem. Res. 34 (1996) 2110–2120. [9] W.P. Walawender, D.A. Hoveland, L.T. Fan, Steam gasification of pure cellulose. 1. Uniform temperature profile, Ind. Eng. Chem. Process Des. Dev. 24 (1985) 813–817. [10] J. Herguido, J. Corella, J. González-Sanz, Steam gasification of lignocellulosic residues in a fluidized bed at a small pilot scale. Effect of the type of feedstock, Ind. Eng. Chem. Res. 31 (1992) 1274–1282.

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Glossary d.a.f. d.a.f.b. FB FBG GY LHV OP SBR SR Tb T⁎b tbed tfreeboard WGSR

dry and ash-free dry and ash-free biomass fluidised bed fluidised bed gasifier (or gasification) Gas yield, Nm3 dry gas free of nitrogen/kg d.a.f.b. low heating value, MJ/Nm3 dry gas oxygen purity, % v/v steam-to-biomass ratio, kg/kg d.a.f.b. stoichiometric ratio, kg/kg mean bed temperature, °C theoretically-calculated mean bed temperature, °C gas residence time in the bed, s gas residence time in the freeboard, s water-gas-shift reaction

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