Fluidized bed gasification of Kingston coal and marine microalgae in a spouted bed reactor

Fluidized bed gasification of Kingston coal and marine microalgae in a spouted bed reactor

chemical engineering research and design 9 1 ( 2 0 1 3 ) 1614–1624 Contents lists available at ScienceDirect Chemical Engineering Research and Desig...

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chemical engineering research and design 9 1 ( 2 0 1 3 ) 1614–1624

Contents lists available at ScienceDirect

Chemical Engineering Research and Design journal homepage: www.elsevier.com/locate/cherd

Fluidized bed gasification of Kingston coal and marine microalgae in a spouted bed reactor Israa K. Alghurabie a,b , Basim O. Hasan b , Brent Jackson a , Adam Kosminski a , Peter J. Ashman a,∗ a b

School of Chemical Engineering, The University of Adelaide, SA 5005, Australia Department of Chemical Engineering, Alnahrain University, Iraq

a b s t r a c t Steam gasification experiments were performed using a low-rank coal from South Australia, a marine microalga, and a blend of leached microalgal biomass and coal, in a spouted, fluidized bed reactor. The effect of different operating conditions – air-to-fuel ratio (A/F), steam-to-fuel ratio (S/F) and bed temperature (Tb ) – on the producer gas composition was investigated. Producer gas compositions were analyzed and samples of bed material were also examined to identify ash components formed during each experiment. The optimum operating conditions for coal gasification, in this system, were identified to occur with A/F = 1.82, S/F = 0.75 and Tb = 850 ◦ C. These conditions resulted in a producer gas with the highest heating value (per mass of fuel fed), the highest extent of carbon conversion and the optimum H2 :CO ratio for Fischer–Tropsch synthesis. In addition, preliminary attempts to gasify a sun-dried marine microalga are reported. The dried biomass, sieved to 1.0–3.35 mm, was gasified with air and steam. Preliminary experiments, utilizing the as-received biomass, proved unsuccessful due to rapid bed sintering. Leaching of the algal biomass to remove the extra-cellular salt and co-gasification of the resultant biomass (10 wt%) with low-rank coal also proved unsuccessful due primarily to blockages of the downstream product lines most likely due to attrition of the algae feed in the screw feeder and elutriation from the bed. © 2013 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved. Keywords: Coal gasification; Microalgae gasification; Agglomeration; Bed sintering; Producer gas; Fluidised bed gasification; Spouted-bed reactor



With increasing cost of crude oil and a rapidly growing global demand for reliable transportation fuels, the use of cheap lowrank coal is seen as one possible pathway for the economical production of synthetic transport fuels. Several companies in Australia are actively pursuing such a technology and previous work at the University of Adelaide has investigated fluidizedbed gasification of certain low-rank coals (McCullough et al., 2011) and the feasibility of using additives to control bed sintering under both gasification and combustion conditions (van Eyk et al., 2012). However, the use of coal in such technologies is prone to environmental concerns, due to the emission of greenhouse gases, and therefore the use of low-carbon alternatives is desirable.

There is a need to identify alternative transport fuels to meet a growing population in the face of dwindling cheap fossil fuel supplies. Biomass has long been considered an attractive alternative energy source due to increasing fossil fuel prices and the worsening effects of global warming (Lim and Alimuddin, 2008). The sustainable production of biofuels from marine microalgae shows enormous potential (Chisti, 2007) however the optimum conversion pathways remain unclear. While renewable alternatives are attractive from an environmental perspective, they are expensive compared to fossil fuels and, in some cases, are not compatible with existing infrastructure, which presents an additional barrier. The New Policies Scenario proposed by the International Energy Agency (IEA) has predicted a world primary demand for energy that will increase by 40% between 2009 and 2035.

Corresponding author. Tel.: +61 8 8313 5072; fax: +61 8 8313 4373. E-mail address: [email protected] (P.J. Ashman). Received 2 January 2013; Received in revised form 23 April 2013; Accepted 29 April 2013 0263-8762/$ – see front matter © 2013 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.cherd.2013.04.024

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This is expected to result in carbon dioxide emissions due to energy production increasing by 20% with a long term rise in global temperature of approximately 3.5 ◦ C (IEA, 2011). These goals are ambitious and require substantial reductions in carbon dioxide emissions through fuel substitution and improvements in energy efficiency. Increasing the efficiency of new coal plants, by the use of advanced technologies, is an important part of the New Policies Scenario, accounting for almost a fifth of the additional required reductions in emissions (Hessen, 1990). Integrated Gasification Combined Cycle (IGCC) is a high efficiency process to generate electricity using solid fuels, such as coal and biomass. Currently it is estimated that transportation consumes around 27% of the world wide primary energy (Uduman et al., 2010), thus reduction in transportation and general engine-based emissions can play a significant role in the reduction of greenhouse emissions. Biofuels such as synthesis gas (syngas), which can be used as a feedstock for the production of synthetic transportation fuels, is a much cleaner alternative to fossil fuels and, if generated at the correct conditions, can result in reduced carbon dioxide emissions. Syngas is produced via gasification, a well-known process, and has been used as a cleaner alternative to combustion for decades (Molina Grima et al., 2002). This process cannot only be used for coal but for other carbonaceous fuels such as biomass with similar results. Marine microalgae have recently received a lot of interest as a new and potentially sustainable biomass source for the production of renewable energy using non-arable land and non-potable water. Microalgae utilize carbon dioxide and nutrients to convert solar energy to chemical energy via photosynthesis (Chisti, 2007). Some of the major characteristics which set microalgae apart from other biomass sources are that algae can have high areal productivities, and thus high biomass yield per unit land area, and can have high lipid content (Chisti, 2007; Rodolfi et al., 2008). Many of the current technologies being pursued focus on the production of lipids, for the manufacture of biodiesel or aviation fuel, with residual microalgae biomass as a byproduct that must be dealt with. One option is to gasify the residual biomass to produce syngas, which can either be used directly for energy production, converted to liquids fuels via Fischer–Tropsch synthesis, or converted to other relatively high-value chemicals. The use of fluidized beds for gasification of biomass is well established (Pereira et al., 2012). However, the gasification of marine microalgae biomass in fluidized beds is expected to be challenging due to its high salt content leading to operational problems such as fouling, and also to agglomeration and defluidisation of the fluidized bed gasifiers. There are several different methods used for the gasification of coal and biomass, including up-draft, down draft, fluidized bed and entrained flow reactors, all with their respective advantages and disadvantages. For this research a spouted, fluidized-bed reactor has been used for the gasification of Kingston coal – a low-rank coal from South Australia – and marine microalgae. Within the spouted bed gasifier there are three specific regions of mixing: a central dilute core (spout), with upward moving solids entrained by a concurrent flow of fluid, a dense phase annular region (annulus), with counter current percolation of fluid, and a “fountain” in which particles are entrained upward from the spout and fall down onto the surface of the bed (Mendes et al., 2008). Spouted bed gasifiers can handle high ash content fuels, provide strong solids mixing and provide a rapid heating rate of the coal in


Fig. 1 – Schematic diagram of the spouted, fluidised-bed reactor. the spout favouring coal devolatilisation and partial gasification (Xiao et al., 2007). This makes the spouted bed reactor ideal for the partial gasification of a wide range of fuels giving extended ranges of flammability, enhanced reaction rates, and super-adiabatic temperatures (Marnasidou et al., 1999). This present study aims to investigate the gasification of Kingston coal and Tetraselmis sp., a marine microalgae, and to examine the effect of varying the gasification operating parameters while firing coal, microalgae and mixtures of these two fuels.




Reactor description

A schematic diagram of the spouted, fluidised-bed reactor is shown in Fig. 1. The reactor has an air feed at the base, lock hoppers at the top for feeding the fuel and a burner for the product gas (producer gas). Air is fed to the reactor at a constant rate of 35 L/min and is preheated using an electrical heater. Water is added to this air at a rate ranging between 0 and 1.2 kg/h and is vaporized in the hot air stream. The steam/air mixture is passed into the base of the reactor and then passes through the bed material (200 g of silica sand) as a jet, forming the so-called spouted-bed. As the air/steam mixture passes through the bed, it interacts with the fuel particles resulting in partial combustion, producing H2 , CO and CH4 , and also heating the reactor due to the heat release associated with the partial combustion. Gases leaving the reactor (producer gas) are sent to the flare burner, which has a pilot-flame fuelled with commercial LPG (Liquefied Petroleum Gas), where it is burned to ensure the almost complete destruction of noxious and toxic compounds, such as hydrogen sulphide. Ash and other particulates are separated using a cyclone and then collected in a separate vessel, which can be removed during the gasification run to facilitate the collection of ash samples. The reactor is stainless steel, comprised of a cone section at the gas inlet which begins with an inside diameter of 12 mm expanding to a cylindrical section with an inside diameter of 77 mm and a height of 1.2 m. A removable cover plate is


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located at the top of the cylindrical section, which allows for the preliminary addition of bed material, with fuel fed to the reactor at the side of the cylindrical section (550 mm above the entrance to the conical base) via a set of lock hoppers and a screw feeder. Nitrogen is supplied to the fuel hopper to create backpressure and maintain an inert atmosphere within each of the hoppers and the fuel feeder. The fuel feed line is also water-cooled at the entrance to the reactor to avoid premature heating of the fuel in the screw feeder. Compressed air is heated using a Leister electric hot air tool type 5000. The hot air is fed into the reactor through a stainless steel mesh (to prevent the back flow of bed material), attached to a pipe which can be removed at the end of a run to retrieve the bed material. The feed water flow rate is measured using a rotameter and fed directly into the hot gas stream where it was vaporized. Heating of the reactor is achieved via four heating elements surrounding the cylindrical section. The heat loss to the surroundings is minimized by ceramic brick insulation, which encloses both the main body of the reactor and the heating elements. Once the reactor reaches a stable temperature, the heating elements are only used to maintain constant reactor temperature by compensating for the heat loss from the reactor since most of the reactor heating is provided by partial combustion of the fuel. Four Type-K thermocouples (3 mm outer diameter) and two pressure tappings are used to measure the reactor temperature at various locations within the bed and the bed pressure drop. The thermocouples are labelled T1 through to T4; thermocouple T1 is located just below the conical distributor, while T2–T4 are located at heights of 35, 65, and 105 mm above the gas inlet, respectively. The pressure tappings are located 15 mm below the gas inlet (P1) and 190 mm above the gas inlet (P2). During stable gasification the temperatures measured are very steady and generally fluctuate by less than ±5 ◦ C.



Low-rank coal from the Kingston deposit in South Australia was used for the coal gasification and co-gasification experiments. The moisture content varied slightly from batch to batch. The coal composition analysis and ash analysis are shown in Table 1. The coal was dried and sieved to a particle diameter of 1.0–3.35 mm for all runs. A sample of marine microalgae, Tetraselmis sp., was obtained from the Muradel Pty Ltd pilot plant located in Karratha. The biomass was grown in 200 m2 raceway ponds, under a range of hypersaline conditions, during commissioning work at the pilot plant. The biomass was harvested and sun-dried before being frozen and shipped to Adelaide. The biomass was thawed and further oven-dried prior to sieving. Two size fractions of dried algal biomass were used in the algae gasification and co-gasification experiments: 1.0–3.35 mm and 1.0–2.0 mm. Silica sand (200 g) was sieved to 0.6 mm diameter and used for the bed material in all gasification and co-gasification experiments.


Experimental procedure


Coal gasification

Two different temperatures were used for these experiments with three different air-to-fuel ratios (A/F) and two different steam-to-fuel ratios (S/F) (see Table 1). For each experiment,

Table 1 – Fuel and standard ash analyses for Kingston coal. Proximate analysis (wt%)


Moisture (as received) Volatile matter (dry basis) Fixed carbon (dry basis) Ash yield (dry basis)

12–18 45.7 38.7 15.6

Ultimate analysis (wt%, daf) C H S Na Cl

56.4 4.3 2.9 0.93 0.16

Ash analysis (wt% in ash) SiO2 Al2 O3 Fe2 O3 TiO2 K2 O MgO Na2 O CaO SO3

33.8 14.2 1.4 0.4 0.5 14.6 5.7 12.4 16.3

the producer gas composition, carbon conversion and gasification efficiencies were determined. Steam-to-fuel ratio (S/F) was calculated (Eq. (1)) accounting for both the water in the coal and the water added as steam. S/F =

Fs + mc · Mc mc (1 − Mc )


where FS is the flow rate of steam into the bed (kg/h), mc is the mass flow rate of coal (dry basis) into the bed (kg/h) and Mc is the moisture content of the coal used (kg H2 O/kg wet coal). The air-to-fuel ratio (A/F) is calculated based on the dry air flow rate and wet coal mass flow rate. In these experiments, the A/F and S/F were varied by altering the coal mass flow rate and the water flow rate, respectively, for a constant air flow rate of 35 L/min. Prior to each run, enough coal for 5 h of operation at the required feed flow rate was dried and sieved. Following the addition of 200 g of sand, the reactor was heated to 450 ◦ C with an air flow of 70 L/min. The air flow was then reduced to 35 L/min and the fuel feed started. After reaching the desired temperature within the reactor, this was maintained using heat input from heating elements surrounding the reactor vessel. Steam was then added at the required flow rate and the experiment was allowed to continue for 4 h, once the conditions reached steady state. After the run was complete, the steam was turned off, the bed allowed to slump and the reactor left to cool. Some runs were completed in less than 4 h due to the occurrence of agglomeration, in which case gas flows were ceased and the slumped bed was left to cool.


Algae gasification

Algae gasification experiments were attempted with a fuel mass flow rate of 1.43 kg/h, steam-to-fuel (S/F) ratio of 0.5 and bed temperature of 850 ◦ C. The steam-to-fuel ratio was calculated using Eq. (1), which accounts for both water in the fuel and water added as steam. Algae biomass was dried and sieved prior to each experiment. An amount of sand (200 g) was added to the reactor prior to the reactor being heated to 450 ◦ C (under the combined action of the external heating elements and the air pre-heater) with

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an air flow rate of 70 L/min. The air flow rate was then reduced to 35 L/min and the fuel feed started. Initially the bed temperature increased to 750 ◦ C. It was found that stable operation was not possible; after 10 min the bed temperature spontaneously decreased to 600–620 ◦ C and then appeared to stabilize. Based on this abnormal operation, it was suspected that defluidisation had occurred and so the gasifier was shut down and left to cool. Visual inspection of the bed material confirmed the presence of agglomerates. Samples of the agglomerate and bed material were analyzed by SEM to determine their structure and composition.


Co-gasification of coal and algae

Co-gasification experiments were conducted, using low-rank coal from the Kingston deposit in South Australia mixed with 10 wt% of marine microalgal biomass (Tetraselmis sp.). For these experiments the gasifier was operated with air-to-fuel (A/F) and steam-to-fuel (S/F) ratios of 2 and 0.5, respectively, with a reactor temperature of 820 ◦ C. Shortly after starting fuel feeding the temperature increased to 820 ◦ C. Gas samples were collected at 30, 45, 60, 75, and 90 min after fuel feeding commenced. The bed temperature continued to rise steadily to 885 ◦ C and the gasifier was then shut down to prevent damage to the reactor. Note that it is a feature of this reactor that during defluidisation events the measured bed temperature increases rapidly.


Gas analysis

Gas samples were collected from the freeboard, using the gas sample valve shown in Fig. 1, into sample bags and then sealed. Each sample was then analyzed used a twin-column Agilent 3000 Micro GC (PoraPLOT Q on channel A and Molecular Sieve 5A with PoraPLOT U pre-column on channel B). Gas samples are preheated in the micro GC to 90 ◦ C prior to column injection. Carrier gases are helium (Channel A) and argon (Channel B). The Cerity NDS for Chemical QA/QC software was used to analyze the spectra and to calculate molar concentrations based on calibrated peak heights. The concentration of nitrogen in the gas sample is used as a tracer to quantify the molar flows of products. In this calculation, the molar flow of fuel nitrogen is negligible as compared to the flow of molecular nitrogen introduced with the air reactant and thus is ignored. The water vapour in the product stream was calculated using an oxygen balance over the gasifier, assuming that all of the oxygen present in the air reactant is consumed within the gasifier.


Results and discussion


Coal gasification

3.1.1. Effect of steam-to-fuel (S/F) ratio on producer gas composition Fig. 2 presents data for the measured concentrations of CO2 , CO, CH4 and H2 in the product gas, as a function of A/F ratio, at various S/F ratios and temperatures. Fig. 2A–C shows the concentrations of CO2 , CO and CH4 , respectively, expressed as a fraction of the solid-phase carbon that is converted into these species. Fig. 2D, on the other hand, plots the concentration (vol%) of H2 in the product gas. At 820 ◦ C, when the S/F ratio is increased from 0.5 to 0.75, the fraction of carbon present as CO2 and CH4 can be seen to increase while fraction of carbon converted to CO decreases. At 850 ◦ C, the experiments give


less conclusive results with weaker trends in the converted carbon. A list of the competing reactions that may occur under gasification conditions (Sharma, 2007; Choi et al., 2001) is presented in Table 3. The trends in producer gas composition with increasing S/F ratio at 820 ◦ C (Fig. 2) may be explained as being due to an increase in the rate of reaction R.3 and a shift to the right in the equilibrium state of reaction R.6 due to an excess of water. An increase in the rate of reaction R.3 should also result in an increase in the CO concentration, however, the excess water present can then react with the CO produced (reaction R.6) resulting in an increase in CO2 and H2 production. The extra H2 produced can also react with CO to produce CH4 (reaction R.7), which explains the observed increase in methane. The opposite is observed when examining the fraction of carbon converted to CO2 and CO for A/F ratios of 1.65 and 2.04. This may be due to an increase in control by endothermic reactions (R.2, R.5–R.8). According to Le Chatelier’s principle, as the temperature is increased, endothermic reactions will increase in activity and exothermic reactions decrease. Hence a higher operating temperature would result in an increased production of CO via reactions R.3 and R.4 which are endothermic. This would also result in less CO2 , as well as less carbon consumption via reaction R.1 which is highly exothermic and responsible for a large amount of the carbon converted to CO2 . The excess H2 and CO present is enough to drive an increase in the production of CH4 via reaction R.8. The steam-to-fuel ratio which gives the highest H2 /CO ratio is 0.75. Air-to-fuel ratios of 1.65, 1.82 and 2.04 give H2 /CO ratios of 2.0, 2.0 and 1.4 respectively at 850 ◦ C. The optimum H2 /CO ratio for fuel and chemicals production depends on the nature of the downstream process. For example, for methanol synthesis the ideal H2 /CO ratio is 2.0. Note that the H2 /CO ratio is not directly observable from the data presented in Fig. 2.

3.1.2. Effect of air-to-fuel (A/F) ratio on producer gas composition The amount of carbon converted to carbon dioxide decreases as the A/F ratio increases from 1.65 to 1.82 in all cases except for one (850 ◦ C, 0.75 S/F) (Fig. 2A). An increase in the oxygen available to react with carbon (A/F = 1.65 to A/F = 1.82) results in an increase in the production of CO2 via reaction R.1. This produces more heat which is quenched in part by reaction R.4 resulting in an increase in the carbon converted to CO and the concentration of the produced H2 . As the oxygen levels are further increased (A/F = 2.04), the amount of carbon converted to CO2 via reaction R.1 increases and the amount of CO converted to CO2 via reaction R.5 increases. This explains the increase in CO2 and the decrease in CO when the A/F ratio increases from 1.82 to 2.04 at 820 ◦ C. With a S/F ratio of 0.5, the amount of steam available for reaction increases and reaction R.3 reduces the amount of carbon available for reaction R.1 resulting in a reduced production of CO2 and increased production of CO (case 2, 820 ◦ C, 0.75, A/F = 1.82 to A/F = 2.04) (Fig. 2A and B). The H2 production increases with A/F ratio at 820 ◦ C and decreases at 850 ◦ C (Fig. 2D). As explained previously, an increase in oxygen in the cases at 820 ◦ C results in a shift to the right in reaction R.6 and more H2 . At increased temperatures (850 ◦ C) the opposite trend occurs with the less exothermic reactions being more favoured. This will result in more CO produced (reaction R.2) as the oxygen concentration increases. This reduces the amount of H2 produced via reaction R.3 and,


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Fig. 2 – Fractional conversion of fuel-carbon to CO2 (A), CO (B) and CH4 (C) in the product gas, and H2 concentration (D) in the product gas (vol%), as a function of bed temperature (T, ◦ C), A/F ratio and S/F ratio. (Legend key: T (◦ C), S/F ratio).

Table 2 – Operating conditions for the coal gasification experiments. Run 1 2 3 4 5 6 7 8 9 10 11 12

Feed flow rate (kg/h) 1.60 1.45 1.29 1.60 1.45 1.29 1.60 1.45 1.29 1.60 1.45 1.29

Air flow rate (l/min)

Average bed temperature (◦ C)

35 35 35 35 35 35 35 35 35 35 35 35

820 820 820 820 820 820 850 850 850 850 850 850

as the availability of oxygen is further increased, the CO2 produced via reaction 1 will increase. An increase in CO2 will shift reaction R.6 to the left further decreasing H2 concentrations and increasing CO. The CH4 levels are dependent mainly on H2 levels in the reactor (reactions R.7 and R.8). The trends observed are similar to those for hydrogen with the exception of the run with

S/F 0.50 0.50 0.50 0.75 0.75 0.75 0.50 0.50 0.50 0.75 0.75 0.75

A/F 1.65 1.82 2.04 1.65 1.82 2.04 1.65 1.82 2.04 1.65 1.82 2.04

A/F = 2.04, S/F = 0.5 and T = 820 ◦ C (Fig. 2C). This run showed mixed results for all components; the CO2 and H2 produced were higher than what would be expected from examining the other runs and the CO lower. The H2 /CO ratio tended to decrease as the A/F ratio increases; at 850 ◦ C the average H2 /CO is 1.97, 1.65 and 1.37, for A/F = 1.65, 1.82 and 2.04, respectively. From these values it

Table 3 – Gas-phase and gas–solid reactions that might occur in the gasifier. Reaction R.1: C + O2 → CO2 R.2: C + 1/2O2 → CO R.3: C + H2 O → CO + H2 R.4: C + CO2 → 2CO R.5: CO + 1/2O2 → CO2 R.6: CO + H2 O → CO2 + H2 R.7: CO + 3H2 → H2 O + CH4 R.8: C + 2H2 → CH4

Reaction name

HR (kJ/mol)

Combustion of carbon Partial combustion of carbon Steam gasification reaction Boudouard reaction Combustion of carbon monoxide Water–gas shift reaction Steam reforming reaction Methanation reaction

−393.5 −110.5 175.3 172.5 −283 2.8 −250.2 −74.9


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Table 4 – Calculated carbon conversion and producer gas heating values for each of the runs shown in Table 2. Run Average carbon conversion Heating value MJ/kg of coal












0.62 6.36

0.54 6.35

0.77 8.99

0.66 7.15

0.75 9.61

0.960 11.72

0.76 9.61

0.77 9.87

0.73 9.22

0.87 11.52

0.96 12.14

12 0.88 10.98

can be seen that a lower A/F ratio yields higher H2 /CO ratios; as A/F ratio is further decreased the molar flow rates of H2 and CO, and the total production of product gas, are also reduced.


Heating value

The heating value of the producer gas, per kg of coal, was calculated and is shown in Table 4 for each of the experimental runs listed in Table 2. The carbon conversion for each run, calculated based on the experimental data, is also shown in Table 4. The heating value, HV (MJ/kg of coal), is calculated using the following equation:

 HV =


(Hc,i × ni )

Fig. 3 – Photo of agglomerates collected after the run at 850 ◦ C, A/F = 1.82, S/F = 0. The agglomerates shown here are approx. 3.5–12 mm in size.



where m is the mass flow rate of coal (dry basis, kg coal/h) and Hc.i is the heat of combustion for species i (MJ/mol), ni is the molar flow rate of species i (mol/h) in the producer gas, and i = C2 H4 , H2 S, C3 H8 , H2 , CH4 and CO. The carbon conversion is calculated from the following equation: Carbon conversion =

mc(gas)out mc,in


where mc(gas)out is the total mass flow rate of carbon in the producer gas (kg/h) and mc,in is the mass flow rate of carbon fed to the gasifier (kg/h). mc(gas)out is calculated using N2 as a tracer by relating the measured concentration of N2 in the producer gas to the known flow of N2 in the air (the flow of N2 to the lock hoppers is negligible and is neglected in this calculation). mc,in is calculated from the known carbon content and feed flow rate of the coal. Producer gas heating values of 11.52, 12.14 and 10.98 MJ/(kg of fuel) for A/F ratios of 1.86, 1.82 and 2.04, respectively, and with S/F = 0.75 and T = 850 ◦ C were the highest obtained in these experiments. These values are strongly correlated to the carbon conversion achieved (Table 4). Both S/F ratio and temperature had a large effect on the conversion achieved. The average carbon conversion tended to increase as the A/F ratio increased due to higher conversion rates achieved as a result of a decrease in the ratio of carbon to oxygen. From this data it would stand to reason that further increasing the A/F ratio would result in higher HVs until a point is reached where the carbon conversion ratio cannot increase any further. An economic analysis over the whole system would need to be done in order to determine the optimum A/F value in terms of cost of energy inputs and value of product gas produced.


Fig. 4 – Micrograph of a SEM secondary electron (SE) image for a section of a typical agglomerate recovered from a coal gasification experiment. point”. Bed particles form sticky surfaces beyond this point. As a consequence particle agglomerates develop resulting in reduced relative movement between particles and ultimately defluidization (McCullough et al., 2011). The agglomerates (Fig. 3) retrieved from the bed material at the end of this run was analyzed by scanning electron microscopy (SEM) using the secondary electron (SE) detector (Fig. 4) and the backscattered electron (BSE) detector (Fig. 5). This gave insight into the typical composition of the agglomerates; two examples can be seen in Tables 5 and 6. The backscattered electron (BSE) detector is used to generate information about the Z contrast of the agglomerate samples (Z = atomic number) such that regions containing atoms with high Z values reflect more electrons and thus are seen as relatively brighter whereas, conversely, atoms with low Z values are seen as relatively dark regions. Samples analyzed using the


Operation was attempted at A/F = 1.82 and T = 850 ◦ C, but in the absence of steam (S/F = 0). This resulted in a gradual but uncontrollable increase in the bed temperature after about 2 h of operation. Once the bed reached a temperature of 900 ◦ C, at which agglomeration was suspected to occur, a sharp increase in temperature to more than 1000 ◦ C was observed. In general, when the bed temperature surpasses a critical temperature agglomeration will occur, this is known as the “sintering

Fig. 5 – Micrograph of a SEM backscattered image (BSE) image for a section of a typical agglomerate recovered from a coal gasification experiment.


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Table 5 – Agglomerate composition measured at the spot marked ‘X’ in Fig. 4.

Table 7 – Agglomerate composition measured at the spot marked ‘X’ in Fig. 6A.





C O Na Mg Al Si K Ca Fe

31.77 21.83 0.58 4.66 1.79 30.26 0.64 7.6 0.85

O Na Mg Al Si S Cl K Ca Fe

10.86 4.97 1 58.19 3.11 1.89 15.17 1.61 2.24 0.97




SE detector were placed directly into the microscope. Thus, the EDS analysis is for a 3D surface and so may be considered as semi-quantitative. Whereas samples analyzed using the BSE detector were coated, sliced and then polished to ensure a flat cross-section for the EDS analysis. Agglomeration during gasification occurs as a result of defluidization. The bed temperature exceeds the “hightemperature defluidization limit” which means a higher superficial velocity is required in order to avoid defluidization for a given bed temperature. Defluidization will initially result in particle growth within the annulus generating a stationary phase. A channel is then generated through the stationary bed material, near to the location of the original spout region, through which air can pass freely. This reactor design does not incorporate any efficient methods for cooling down in the event of excess heat generation. This made it difficult to control temperature under such conditions. Adding additional steam or reducing the air flow rate can be used to alleviate this however for experimental purposes these two variables remained constant during gasification runs. Without sufficient levels of steam present in the reactor reactions R.3 and R.7 were unable to act to consume some of the excess heat.


Microalgae gasification

Gasification of marine microalgae was attempted for a fuel mass rate of 1.43 kg/h, steam-to-fuel (S/F) ratio of 0.5 and bed temperature of 850 ◦ C. Immediately upon commencement of fuel feeding to the preheated (450 ◦ C) reactor, the bed temperature increased to 750 ◦ C and then, after about 10 min, was observed to decrease to between 620 and 600 ◦ C and stabilize. Based on this abnormal operation, it was suspected that defluidisation had and so the gasifier was shut down and left to cool. The bed was found to have agglomerated in this case. Samples of the agglomerates were analyzed using SEM and typical

Table 6 – Agglomerate composition measured at the spot marked ‘X’ in Fig. 5. Element


O Na Mg Al Si S Cl K Ca

47.66 0.35 0.35 0.74 49.28 0.39 0.4 0.22 0.6




Table 8 – Agglomerate composition measured at the spot marked ‘X’ in Fig. 6B. Element


O Na Mg Al Si P S Cl K Ca Fe

16.88 15.97 7.42 6.12 2.06 1.3 2.84 37.38 3.97 2.76 3.29



images are shown in Fig. 6. The composition of the agglomerate at the points marked ‘X’ in Fig. 6A and B is reported in Tables 7 and 8, respectively. Images of agglomerates were also collected using the SEM in backscattered electron (BSE) mode and these are shown in Fig. 7 with the composition at the various locations in these figures reported in Tables 9 and 10. A sample of the bed material for this run was also analyzed by SEM-EDS to determine the composition of this material. SEM images of the bed material are shown in Fig. 8 with the composition of the sample at the marked points reported in Table 11. These data show relatively large fractions of Na, Al, Mg, and Cl ions present in the examined samples. These would most likely be chloride salts due to the seawater media used to grow the algae. The salt is likely to have crystallized during the process of drying the algae. Relatively pure silica oxide is

Table 9 – Agglomerate composition measured at the spots marked ‘1 , ‘2 , and ‘3 in Fig. 7A. Element O Na Mg Al Si P S Cl K Ca Total

Spot ‘1’ wt%

Spot ‘2’ wt%

Spot ‘3’ wt%

48.97 0 0.41 1.09 47.45 0 0.41 0.46 0.42 0.79

44.41 2.85 4.32 14.91 20.8 0 0.32 1.51 2.69 8.2

59.5 4.13 6.85 9.93 3.32 4.23 3.89 5.82 0.95 1.38




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Fig. 6 – Micrographs of SEM secondary electron images for two sections of the agglomerated bed material recovered after feeding microalgae to the gasifier.

Fig. 7 – Micrographs of SEM backscattered images for two sections of the agglomerated bed material recovered after feeding microalgae to the gasifier.

also observed in the samples, which is due to the sand used for the bed material. To test for the level of salt present in the algae a sample was immersed in de-mineralized water (conductivity = 0) and agitated. After 3 h the conductivity of the microalgae solution was measured giving a value of 13.3 mS. The water was replaced three times to ensure that all of the salt was dissolved and then the solution was left to settle, the supernatant removed and the remaining material left to air dry. The algae was then put into the oven to further dry to a moisture content of less than 20% (for further analysis and for use in the co-gasification experiments). The dried algae was analyzed

using SEM with the secondary electron (SE) detector, both before leaching (Fig. 9A) and after leaching (Fig. 9B). The corresponding spot analyses (Tables 12 and 13, respectively) show that the NaCl content, relative to the other elements present, decreased substantially due to leaching. A basic analysis was completed to calculate the volatile, salt, and ash components in algae. This was done by placing a known weight of dried leached algae into a furnace at 840 ◦ C for 3 h. This combusted all volatile matter leaving behind any non-volatiles, the crucible was then taken out of the oven and weighed. The sample was then added to de-mineralized water and the conductivity and salinity were measured to calculate the salt content of the

Fig. 8 – Micrographs of SEM secondary electron images for two sections of the agglomerated bed material recovered after feeding microalgae to the gasifier.


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Fig. 9 – Micrographs of SEM secondary electron images of the raw algae before leaching (A) and after leaching (B).

Table 10 – Agglomerate composition measured at the spots marked ‘1 and ‘2 in Fig. 7B.

Table 13 – Composition data for raw algae after leaching, as measured using SEM at the spot marked ‘X’ in Fig. 9B.




O Na Mg Al Si P S Cl K Ca

45.8 3.01 6.38 10.21 1.41 1.91 12.26 8 1.55 9.48

O F Na Mg Al Si P S Cl K Ca Total

Spot ‘1’ wt%

Spot ‘2’ wt%

2.12 0 33.74 0 0.7 0.73 0.82 1.75 58.56 0.79 0.79

40.01 0 3.08 20.39 6.39 20.66 0.89 0.67 4.24 1.17 2.51


Table 11 – Agglomerate composition of the bed material samples measured at the spots marked ‘X’ in Fig. 8A and B, respectively. Element O Na Mg Al Si P S Cl K Ca Fe Total

Fig. 8A (wt%)

Fig. 8B (wt%)

39.05 4.14 24.53 0.79 18.88 0 1.57 7.8 0.29 1.49 1.45

5.91 30.8 0 1.04 0.55 0.51 1.96 54.15 1.01 2.85 1.21






non-volatile matter (which was found to be 3.59%). From this the ash content of the raw microalgae algae was found to be approximately 38%.


Co-gasification of coal and leached algae

The leached alga was co-gasified with Kingston coal in the ratio of 10 wt% algae to 90 wt% coal. After commencing fuel feeding the temperature increased rapidly to 820 ◦ C, successive gas samples were drawn after 30, 45, 60, 75 and 90 min and at these times the bed temperature was measured to have increased to 830, 840, 860, 870 and 880 ◦ C, respectively. As the bed temperature continued to rise steadily, the gasifier had to be shut down in order to avoid damage of the reactor due to excessive temperature. Fig. 10 reports the mole fraction of

Table 12 – Composition data for raw algae before leaching, as measured using SEM at the spots marked ‘1’ and ‘2’ in Fig. 9A. Element O F Na Mg Al Si P S Cl K Ca Total

Section 1 wt% 1.99 0.35 39 0 0.49 0.53 0.61 0.84 55.02 0.6 0.58 100

Section 2 wt% 4.36 0.12 37.79 0 0.73 0.61 0.12 0.52 55.61 0 0.15 100

Fig. 10 – Producer gas composition (vol%) measured during the co-gasification experiment. Gas samples were collected after specific time intervals, which corresponded to the bed temperatures shown (see text for details).

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Fig. 11 – Photos of the outlet tubes of the reactor vessel showing the large extent of deposition due to a grey/black substance. H2 , CO2 , CH4 and CO in the producer gas, as a function of bed temperature, during this test. While agglomeration was not observed to occur, the bed pressure drop and bed temperature increased to high values of approx. 500 Pa and more than 900 ◦ C, respectively, after 90 min. This might have occurred due to decreased bed fluidization, or it could have been due to blockages in the gasifier downstream of the reactor which prevented part of the produced gas from leaving the reactor, thus causing high temperature and pressure inside the reactor. An alternative possibility, that pyrolysis of the algae fuel may have occurred within the fuel feeding system, was unlikely since the screw feeder is water-cooled specifically to ensure that the feed is not heated prematurely. After the reactor was cooled, the bed material was recovered and no agglomerates were found. The product gas outlet pipe (downstream of the reactor vessel) was disconnected and found to be blocked with a large amount of a grey/black substance, possibly ash mixed with unburned fuel and/or tar (Fig. 11). A sample of the grey material blocking the pipes was taken and tested for salt content. The salt level in the blockage was the same as in the raw algae which suggests that the blockage could be attributed to the presence of raw algae. It is expected that the dried algae, which was relatively soft and friable, may have been crushed in the screw feeder and thus entered the reactor as a fine dust, rather than as discrete particles of size 1.0–3.35 mm (as fed from the fuel hopper). In which case, the fine dust would have been elutriated from the bed and carried over into the gas outlet of the gasifier where it became trapped. This would have led to a high pressure drop across the gasifier and ultimately to failure of the bed to fluidize.



The effect of temperature, A/F ratio and S/F ratio on the composition of producer gas generated in a 77 mm spouted-bed fluidised bed gasifier was observed. The equilibrium dynamics varied with temperature resulting in endothermic reactions gaining preference at higher temperatures. This resulted in better gasification as well as higher carbon conversions. As the S/F ratio was increased the hydrogen, carbon dioxide and methane production increased and carbon monoxide decreased. This resulted in reduced carbon conversions and lower heating values of the producer gas generated. The A/F ratio tended to give a similar trend with higher carbon conversions achieved at higher A/F ratios resulting in higher heating values. This was the general trend, it should however be noted that operation at 850 ◦ C resulted in maximum carbon conversion and heating values at an A/F ratio of 1.82. Operation with A/F ratio = 1.82, S/F ratio = 0.75 at 850 ◦ C generated the producer gas containing the highest heating value per mass of coal fed


into the gasifier. This run gave the highest carbon conversion ratio, as well as a H2 /CO ratio of 2.0. This gasifier was unable to operate at low S/F ratios due to issues with agglomeration. This data still gives a good insight into the behaviour of Kingston coal under fluidized bed conditions. Agglomeration of the bed occurred with dried algae as the feedstock, apparently due to the high salt content of the raw algae. To avoid this leaching of the raw algae was performed to reduce the salt content. Co-gasification of coal and the leached algae was then performed. The temperature of the bed was unstable and increased rapidly with a high pressure drop across the gasifier due to blockages of the downstream gas outlet which caused by the elutriation of biomass which had been crushed in the screw feeder. In future work, a different feeding strategy will be necessary to avoid these issues. Further work is underway to resolve these operational problems.

Acknowledgements This research was supported under Australian Research Council’s Linkage Projects funding scheme (project number LP100200616) with our industry partner SQC Pty Ltd. Ms. Alghurabie acknowledges financial support from the Ministry of Higher Education and Scientific Research, Iraq. Mr. Jackson was supported by a University of Adelaide Summer Research Scholarship. The authors acknowledge the contributions of Dr. Philip van Eyk, for his valuable advice and feedback, and the staff in the School of Chemical Engineering Workshop, for their assistance with maintenance of the spouted-bed gasifier.

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