Global energy demands are increasing and so too is the need for more renewable and sustainable sources of energy to help transition us to a post-fossil-fuel-powered world. The European Union (EU) has recently increased its renewable energy target to 32% for 2030 (1), with many countries planning to ban internal combustion engine powered cars by 2040 or sooner. However, the transportation industry is one of the most challenging sectors to adapt to using low-carbon fuels. Transportation modes such as aircraft, heavy-duty and marine vehicles demand high power and energy capacity that are currently unmet by renewable technologies. In the interim, we need clean, sustainable methods, continuous improvement and new innovations in renewable fuels to meet EU and other similar worldwide targets.

Johnson Matthey and bp have been collaborating for the past two decades (2, 3) to develop an efficient reactor system and catalyst for the FT process. This offers a cost-effective method of converting any carbon source into high-quality liquid hydrocarbon fuels.

Today the world consumes more than 55 million barrels (bbl) day^–1 of transportation fuels (4), the vast majority of which originates from crude oil. In addition to being a finite resource, each barrel of crude-oil-derived fuel typically contributes about 475 kg of CO₂ into the atmosphere over its life cycle (based on a 2010 European average) (5). At the same time, hundreds of millions of tonnes of municipal solid waste (MSW) are incinerated or sent to landfill each year, while similar quantities of woody biomass decompose to CO₂ and methane (a more potent greenhouse gas than CO₂) (6). Industrial processes release more than 8 billion tonnes of direct CO₂ emissions (7), and flaring of natural gas releases a further 275 million tonnes of CO₂ equivalent per year (8).

These wastes and emissions are rich in carbon which can be extracted through gasification, reforming or capture of CO₂ (which can then be converted to synthesis gas (syngas) via reverse water gas shift). Converting this carbon to useful syngas rather than allowing it to be emitted to the atmosphere as CO₂ creates an opportunity to significantly reduce fuel life-cycle emissions and its impact on global warming. The carbon intensity of the resulting fuel can typically be reduced by more than 70% (relative to conventional fuel), with reductions of more than 100% possible as grid electricity becomes increasingly renewable and CO₂ sequestration is added to the technology mix.

Gasification of biomass is not new technology. However, it has typically been used to produce syngas for the generation of electricity, rather than chemical synthesis. Effective FT synthesis requires a specific ratio of hydrogen and carbon monoxide, and the catalyst used for this reaction is particularly susceptible to poisoning by impurities that may be present. Gasification of waste introduces the potential for a wide range of contaminants that need to be removed.

The range of potential poisons, and the low levels necessary to ensure continued high-level performance for the FT synthesis catalyst, make syngas purification a critical step of the process. Effective removal of these impurities also ensures that FT products are ultra-clean and high purity (9).

As a world leader in purification and pre- and post-treatment of syngas gas for downstream applications, Johnson Matthey has a range of solutions to condition the syngas (irrespective of its original source) ready for conversion into sustainable fuel.

The FT process was originally developed by Franz Fischer and Hans Tropsch in 1925. It is a way of converting any carbon source into liquid hydrocarbon via syngas, effectively creating synthetic fuel (see Equation (i)).

(i)

Syngas can be generated from various carbon sources, including coal, natural gas, MSW and biomass. The process mainly produces linear, long-chain paraffins that require further upgrading to produce liquid fuels, such as diesel and kerosene. The upgrading step comprises catalytic hydrocracking to both isomerise and crack the long-chain paraffins into smaller-chain paraffins with the correct properties for fuel applications. The various stages in the process are shown in Figure 1. In the quest for sustainable fuel solutions, FT-derived synthetic fuels provide a cleaner way to power cars, heavy-duty vehicles and aeroplanes. The latest developments in FT technology mean that the production of fuel from sustainable carbon sources is now closer to being commercially viable at all industrial scales.

Fig. 1

Catalysts are required for the FT process to increase the rate of reaction and make the process industrially viable. There are broadly two options for FT synthesis, using cobalt or iron catalysts (10, 11). While cobalt is more expensive than iron, it mainly produces normal paraffins. Iron catalysed FT synthesis also incorporates the water-gas shift reaction for CO₂ products and makes a mixture of olefins and paraffins.

The most commonly used catalyst is cobalt, due to its high activity, selectivity to liquid hydrocarbons and stability. Commercial synthesis of hydrocarbons occurs at moderate temperatures of 200–240ºC and pressures of 20–40 bara. During the process, hydrogen and carbon monoxide are converted into long-chain paraffins or waxes over the supported catalyst (see Figure 1). Pore diffusion and mass-transfer effects therefore play a key role in FT catalyst performance due to the need for hydrogen and carbon monoxide to move into and along the catalyst pores against the movement of product molecules going the other way.

The FT synthesis reactions are all highly exothermic, making efficient removal of heat essential for any reactor design. There are a number of benefits of using conventional fixed-bed tubular reactors (12, 13) which is why Johnson Matthey and bp have favoured this design. They are a proven technology with many manufacturers able to fabricate reactors at large scale. They work by holding the catalyst in place via a static bed, which has the advantage of preventing catalyst loss, which could lead to product contamination as can occur in slurry reactors. The reactors have a modular design, which makes increasing capacity as simple as adding tubes; however, conventional fixed-bed tubular reactors are limited by the need to balance tube diameter and catalyst pellet size to achieve effective temperature control without excessive pressure drop. These reactors generally contain tens of thousands of tubes of around 25 mm diameter, resulting in high construction costs with catalyst pellets in the range of 1–2 mm diameter, which reduces catalyst productivity and selectivity to hydrocarbon liquids.

An alternative synthesis route is through slurry reactors. This type of reactor is more efficient at heat removal and uses catalyst powder of the order of tens of microns diameter to minimise pore diffusion resistance. However, slurry reactors can suffer from catalyst attrition, which leads to catalyst loss and product purity issues, and are also less straightforward to scale up compared to fixed-bed alternatives.

Since 1996, Johnson Matthey and bp have been collaborating to bring FT synthesis to the industrial scale. The first major joint venture in 2002 was to build the Nikiski demonstration plant in Alaska, USA (Figure 2), based on the first generation (Gen1) FT catalyst contained within conventional tubular reactor technology (14). The Nikiski plant produced a nominal 300 bbl day^–1 of synthetic crude product from pipeline natural-gas feedstock; and by the time it was decommissioned in 2009, the plant had exceeded all its performance goals related to catalyst productivity, hydrocarbon selectivity, carbon monoxide conversion, methane selectivity and catalyst lifetime. A single charge of catalyst ran for just over 7000 h enabling Johnson Matthey to predict an expected three-year lifetime without any regeneration.

Fig. 2

The integrated plant combined three processes for testing FT technology: a novel compact reformer for syngas generation; a fixed-bed FT reactor; and mild hydrocracking of FT waxes to produce synthetic crude. The original fixed-bed tubular reactor technology was developed as a method of monetising stranded natural gas in remote locations. However, it was only competitive at large scale, above 30,000 bbl day^–1 (~3850 metric tonnes per day (mtpd)), in areas with low natural gas prices and high oil prices.

More recent interest in FT technology is in small-scale applications to produce renewable fuel from MSW or cellulosic biomass. This involved developing technology that lowered costs whilst improving efficiency. In 2009, Johnson Matthey designed a novel catalyst carrier device to fit inside a tubular reactor that allows for the use of smaller catalyst particles. At the same time, bp developed an improved second generation (Gen2) catalyst formulation (15). Both organisations then worked to combine both the new catalyst and the novel catalyst carrier device, which produced a step change in commercial FT performance (see Figure 3). The CANS^TM catalyst carrier technology received global recognition, winning both the Research Project Award and the Oil and Gas Award at the Institution of Chemical Engineers (IChemE) Global Awards in 2017, and the Rushlight Clean Energy Award and Rushlight Bioenergy Award in January 2020. These accolades demonstrate how advanced FT technology will dramatically impact the chemical engineering industry, with many real-world applications.

Fig. 3

The novel catalyst carrier reactor design (16) combines the advantages of the fixed-bed tubular reactors and the slurry-phase systems. Its modular design enables low-risk scale-up and simple operation, while the smaller catalyst particles offer high productivity and selectivity. The stacked catalyst carriers have a unique design that aids their ability to perform FT synthesis as shown in Figure 4.

Fig. 4

A reactor tube contains 60–80 of the CANS^TM catalyst carriers and effectively creates a series of mini adiabatic radial-flow reactors with interbed cooling. Radial flow through each CANS^TM catalyst carrier means that, although the reactor tubes are 10–15 m long, the effective catalyst bed thickness is only around 15% of the overall tube length. This enables the use of sub-millimetre catalyst particles, which improves selectivity and activity whilst limiting the reactor pressure drop to that of a conventional fixed-bed tubular reactor. Wide-diameter tubes of 75–100 mm are used in the novel reactor, which has the effect of reducing the heat-transfer surface per unit volume of catalyst. However, this is compensated by a larger temperature difference at the wall, where reactants are hottest (as opposed to the centre of the tube in conventional fixed-bed tubular reactors). Combining this structure with a high gas velocity through the narrow annulus between the CANS^TM catalyst carrier body and tube wall results in excellent heat transfer. By separating heat removal from the catalyst bed, good control of the reaction temperature is also achieved without the risk of quenching the reaction. The advanced reactor technology also enables operation with >50% inerts in the reacting gas, allowing a single-stage FT reactor to be used in a recycle loop to maximise overall conversion of carbon monoxide to >90% (Figure 5).

Fig. 5

Compared to conventional fixed-bed tubular reactors, the new CANS^TM catalyst carrier and optimised catalyst reduces the number of reactor tubes by 95%, significantly simplifying the design and fabrication of the reactor, resulting in a reduction of capital expenditure costs of around 50% for the FT unit. There is also a three-fold increase in production for the same size reactor as the catalyst performance is closer to that of a powder, with excellent heat and mass transfer to, from and within catalyst particles. The increased productivity at least halves the catalyst volumes usually required for the same production rate. Additionally, containing the catalyst inside the CANS^TM catalyst carrier removes the requirement to filter the catalyst from the wax product. Instead the catalyst is easily replaced by removing the entire catalyst carrier, meaning there is no interaction with the hazardous cobalt catalyst material. Fundamentally, this makes FT applications possible at both small and large scales, with around 6000 bbl day^–1 (770 mtpd) achievable in a single reactor of around 900 tonnes. For areas with tighter transport restrictions, 2000 bbl day^–1 can be delivered in a single reactor of around 4 m diameter and 250 tonnes in weight.

One of the main challenges to overcome was proving that the concept worked at commercial scale, and so Johnson Matthey has invested in extensive testing to develop the engineering science necessary to implement the novel reactor concept. This involved building customised rigs to validate heat-transfer performance, hydraulics and accurately measure reaction kinetics on high-throughput microreactors. Significant engineering effort has been required to develop models capable of accurately predicting the performance of commercial-scale reactors. The initial proof of concept work was carried out using CANS^TM catalyst carriers manufactured by an experienced prototyper.

FT catalysis is strongly influenced by cobalt crystallite size, support properties and catalyst treatments. Selection of the cobalt crystallite size is critical to obtaining the required performance. Larger cobalt crystallites result in a less active catalyst due to the lower surface area to volume ratio, and thereby require higher temperatures to achieve a target conversion. Alternatively, if cobalt crystallites become too small the chemistry favours chain termination (methane formation) over chain growth (C–C coupling). While the desirable cobalt crystallite (17, 18) size is in the narrow range of 8 nm to 10 nm for FT synthesis, the activity of a good catalyst can be significantly reduced by suboptimal treatments, such as the reduction stage of the cobalt oxide to the active metallic phase. Cobalt-based FT catalysts are normally made by impregnating the support with a cobalt salt, calcination to give cobalt oxide and subsequent reduction under hydrogen in the plant to give the active cobalt metal phase. The catalyst reduction process is defined in Equation (ii), which highlights the significant levels of water that are produced throughout the catalyst bed, and this in turn can sinter, reoxidise or damage the catalyst significantly if not fully catered for under process conditions. Catalyst bed profile effects are also significant as the bottom sections of the bed are exposed to the water produced at the top of the bed, while higher pressures required commercially also lead to higher water partial pressures in the catalyst pores.

(ii)

bp originally developed the Gen2 catalyst formulation as a drop-in for conventional fixed-bed tubular reactors and this formulation has been adapted to the CANS^TM technology by Johnson Matthey to produce sub-millimetre size catalyst particles at scale. Developing the new formulation to achieve improved activity, selectivity and stability, while optimising the catalyst activation has required thousands of hours of testing at laboratory scale in high throughput and pilot plant test units. This has been supported by a state-of-the-art FT unit, with exceptional online analytics for all products up to C18 and analytical capabilities such as in situ X-ray diffraction and temperature programmed reductions which enable catalyst evaluation under process conditions (19, 20).

While the hydrocarbons and oxygenates that were identified are known compounds formed during the low temperature, cobalt catalysed, FT process the combination of the multiple analysis techniques used has allowed a level of detail to be gained on the FT product composition that is seldom reported (9). Typically, the long-chain 1-alcohols and carboxylic acids were found to be present at levels of one tenth and one thousandth that of hydrocarbons of equivalent carbon chain length respectively. Additionally, hydrogen-1 nuclear magnetic resonance (¹H-NMR) and carbon-13 nuclear magnetic resonance (¹³C-NMR) analyses were used to quantify the average class compounds concentration of 1-olefin, cis- and trans-2-olefins, 1-alcohol and aldehyde as appropriate for the technique used. The 1-olefin:n-paraffin ratio in the hydrocarbon liquid and wax products was found to decrease significantly with increasing carbon chain length in both phases and much more so than those of the 2-olefin or 1-alcohol.

Catalyst activity and selectivity is only part of the process however, with stability, robustness to process events and life duration also playing a vital role in a commercial catalyst. Johnson Matthey Davy and bp built on their extensive experiences of the Gen1 catalyst in the Nikiski demonstration plant to optimise this further for the Gen2 catalyst. This included several catalyst life tests which operated for many thousands of hours at steady FT process conditions. This included the catalyst formulation used in CANS^TM catalyst carriers operating with exceptional performance over an 18,000 h life test. The gradual drop in catalyst activity over this period was compensated by an increase in operating temperature within the reasonable limits of a commercial reactor. Despite frequent shutdowns and other challenges associated with laboratory-scale operation, the catalyst was still showing good activity and selectivity at the end of this test. This is a result of the process having been designed to be robust and operate in chemically stable conditions.

The CANS^TM catalyst carrier concept has been successfully demonstrated at commercial scale on a pilot plant at Johnson Matthey’s research and development (R&D) facilities in Stockton-on-Tees, UK. It is not practical to test a full-length commercial reactor tube at these facilities, due to limitations on gas supply and product storage capacity, so a creative approach was required. Flexible design of the pilot plant enabled testing of multiple commercial-size CANS^TM catalyst carriers in a much shorter tube; by recycling gas, liquid products and produced water to simulate the full range of conditions and flowrates present in a commercial reactor tube. This, coupled with raising steam in the reactor cooling jacket, has enabled full demonstration of the catalyst, CANS^TM catalyst carriers, hydraulics and heat transfer at commercial conditions, flows and tube diameters.

Over 20,000 h of testing under commercial flowsheet conditions has demonstrated the performance of the CANS^TM catalyst carrier and the Gen2 catalyst with a confirmed product slate and stable catalyst life. A C5+ selectivity of around 90% and C5+ productivities in excess of 300 g l^–1 h^–1 have been demonstrated on the pilot plant. The crude FT product consists of a wax stream which is liquid at reaction conditions and solid at ambient temperature and a light hydrocarbon condensate stream which is liquid at ambient temperature. Figure 6 shows both these products are high quality, clean and catalyst-free.

Fig. 6

The International Energy Agency, France, has measured the share of global energy-related CO₂ emissions from transport at 23%, with aviation contributing 2–3% of worldwide anthropogenic CO₂ emissions (21). There is great potential for this figure to be reduced by using synthetic fuels from sustainable feedstocks, and this makes fuels produced via the FT process an attractive alternative to current aviation fuels. Synthetic fuels also burn cleaner, due to the absence of sulfur and aromatics, while also producing fewer particulates (22). As a result, FT fuels lead to increased combustion and turbine life, while the enhanced thermal stability reduces deposits on engine components and fuel lines. This results in a dual advantage for the aviation industry, in terms of both improved fuel economy and less maintenance of aviation equipment.

The scale of the market is substantial, with air transportation alone expected to consume at least 500 million tonnes per year (11 million bbl day^–1) of fuel by 2050 (23). Practical limitations on the supply of waste feedstocks or local, low-cost renewable power for hydrogen production typically limit the scale of each project to less than 5000 bbl day^–1. With tens of thousands of filled CANS^TM catalyst carriers required for each project, the ability to consistently and efficiently produce both the CANS^TM catalyst carriers and the FT catalyst which they contain is crucial for successful commercial deployment of the technology.

To address this, Johnson Matthey has collaborated closely with a company skilled in delivering sustainable engineered solutions for vehicle exhaust after treatment systems, to develop a mechanical design for the CANS^TM catalyst carriers that is economical to make, can be easily filled with catalyst and meets the functional specifications developed by Johnson Matthey. This has been confirmed by testing of commercial prototypes.

A production line has been constructed and commissioned for mass manufacture and catalyst filling of the CANS^TM catalyst carriers, and the first charge has now been produced for the first commercial project. The production line is fully automated to allow the safe filling of the cobalt-containing catalyst and contains state of the art equipment and in-line quality control to assure the CANS^TM catalyst carriers meet the required functional specifications. The functional specifications were established during the development of the CANS^TM catalyst carriers from concept to prototype with testing performed on in-house built rigs at Johnson Matthey in Teesside, UK. Identifying these upfront allowed Johnson Matthey and the manufacturer to work together to ensure the resulting production line would safely, efficiently and consistently produce high-quality units.

A collaborative team of engineers from Johnson Matthey and the manufacturer worked closely during initial commissioning of the line to work through the various challenges associated with scale-up to mass manufacture of a novel process. This knowledge will be invaluable as further improvements and optimisations are implemented.

In order to achieve high-quality performance, challenging activity and selectivity targets were set for the FT catalyst. Scale-up of the chosen formulation took place at Billingham in the UK. Catalyst preparation was initially at laboratory scale with short-term and long-term testing of development samples conducted using both micro-reactors and CANS^TM catalyst carriers, facilitating accurate modelling of the performance of a full-scale FT reactor.

As scale-up continued, preparation of the FT catalyst moved into the Manufacturing Science Centre (MSC) where appropriate technologies were identified for each of the steps involved in catalyst production. The technical risk of scale-up was minimised by using down-scaled versions of full-scale production equipment.

A fully developed catalyst manufacturing process was transferred from the MSC to a dedicated production asset located at Clitheroe in the UK (Figure 7). Careful attention was paid to the specification of the raw materials used. To ensure a proper understanding of the impact of trace impurities on long term FT catalyst performance, a series of experiments were conducted in which the FT catalyst was doped with different FT poisons.

Fig. 7

Every production batch of FT catalyst has been tested against an agreed quality assurance specification. Conforming product was loaded into CANS^TM catalyst carriers as it was manufactured, thus minimising the overall production timeline.

The Johnson Matthey Davy/bp FT technology incorporating CANS^TM catalyst carriers offers benefits to both small- and large-scale operations with good economics, opening up the prospect of exciting future applications.

Fulcrum BioEnergy, USA, is the first to licence the Johnson Matthey Davy/bp FT technology in its Sierra BioFuels Plant, located near Reno, Nevada (Figure 8). The Sierra plant will be the first in the USA to produce a renewable low-carbon transportation fuel from MSW or household garbage. The plant will first sort the waste to recover recyclables and remove material not suitable for processing, so is not in competition with recycling processes. The remaining material will be processed into a feedstock before being fed into a gasification system to produce a syngas. This is then converted into hydrocarbons by the FT technology for the production of renewable fuels. The Sierra plant construction is approaching completion and when operational will convert approximately 175,000 tonnes of MSW into approximately 42 million litres of renewable FT product each year.

Fig. 8

Multiple projects are being developed in the USA and Europe, which can make a significant contribution to meeting the demand for renewable transportation fuels in the next decade.

In order to meet greenhouse gas emissions reduction targets, especially for aviation (24), production of sustainable fuels will have to substantially increase. There are a range of sustainable fuels potentially available, but limitations on sustainable feedstocks and viable technology routes mean that diesel and jet fuel production via FT synthesis will need to form a key part of this industry. Producing fuels via FT synthesis is not new. However, cost of production was always a barrier, with existing large-scale producers of FT fuels unable to economically scale down to match the size of the waste facilities that feed them. The CANS^TM technology addresses this problem, offering an economic and efficient solution at the scales required by the industry.

The proposed solvent plant is co-located with a Kraft paper and pulp mill in China with throughput as defined in Table I. Figure 1 outlines the Kraft process, which conventionally directs weak black liquor to multi-effect evaporators, producing strong black liquor which is combusted in a Tomlinson boiler to produce steam (4). This steam makes the mill self-sufficient in steam and electrical energy. Importantly, the cooking chemicals (NaOH and Na₂S) are recovered and recycled to the pulping process.

Fig. 1

As previously mentioned, investments in heat integration have freed up a portion of the weak black liquor coproduct for alternative uses. This study explores the opportunity of utilising this excess coproduct, taken as 25% of total production, for isopropanol and acetone production through aerobic fermentation in an integrated solvent plant as outlined in Figure 1.

Given black liquor has no economic value as a product, it is costed at its utility value. This is calculated based on its conventional use for renewable electricity generation, requiring capital investment in increased steam turbine capacity. The foregone net present value (NPV) associated with this conventional use is used as the utility value for the black liquor feedstock.

In the proposed solvent plant (Figure 1), weak liquor undergoes SCWG to carbon dioxide and hydrogen. A challenge, however, is the efficient recovery of the cooking chemicals from the SCWG reactor and their recycle to the pulp mill digestor. Loss of these salts would result in a significant cost to the pulp mill. Under supercritical conditions, the properties of water change from polar to apolar, where the solubility of inorganic salts is very low (31). Cao et al. described the precipitation of alkali sodium salts in SCWG, reporting a neutral pH for the reactor effluent, suggesting that under supercritical conditions the salts largely precipitate from the solution (32). However, this precipitation can cause issues with plugging and fouling within the reactor (33). In this study the salts are removed prior to entering the SCWG reactor, in a manner similar to supercritical water desalination (34, 35) and modelled for SCWG of black liquor (33).

The solvent plant’s mass and energy balance was informed by experimental data from continuous gas fermentation (28), and rigorous process simulation using Aspen HYSYS v11. The lignin content in black liquor was modelled as guaiacol, a model compound for lignin (36), as principal feed to the solvent plant. The weak black liquor is further diluted prior to entering the SCWG reactor, as lower biomass concentrations promote superior thermal cracking and yields greater hydrogen and carbon dioxide owed to the increased water concentration favouring the forward water-gas shift reaction (37).

The simplified flow diagram (Figure 1) outlines the six plant sections of the solvent plant, whilst Figure 2 presents a detailed process flow diagram and operating conditions for upstream and downstream processing. The unit operations included in each of the six plant sections are summarised in Table II. Table III summarises the scale-up of the experimental gas fermentation data for the process simulation, which recognises the oxygen mass transfer limitations associated with the safety requirement to maintain non-flammable operating conditions. The heat integration between the low temperature exothermic gas fermentation and the high temperature endothermic SCWG is facilitated using a heat pump with isopentane as the working fluid (28).

Fig. 2

Isopropanol and acetone are produced in both the aqueous and vapour phase of the bioreactors. The solvents in the vapour phase are recovered via gas absorption through mass integration using internal process streams, i.e. the isopropanol product was utilised to recover acetone, and water to recover isopropanol. For the isopropanol in the aqueous phase, azeotropic distillation is required due to the homogeneous minimum boiling point azeotrope formed between isopropanol and water (38). Conventionally, this azeotrope is broken using an entrainer, historically benzene (39). However, owed to its carcinogenic properties, alternative entrainers such as cyclohexane have been adopted (40). An alternative azeotropic separation technique is pressure swing distillation, taking advantage of the composition differences in the azeotrope at different pressures (41). In this work, pressure swing distillation was employed with the coproduct acetone acting as an unconventional entrainer. Further detail of the separation train is presented in Figure 2.

A U-loop bioreactor, similar to the one used by Peterson et al., is used in this work (42). The benefit of a U-loop bioreactor is that high mass transfer coefficients can be achieved without the need for mechanical agitation, leading to greater oxygen transfer rate and a reduced power requirement compared to conventional stirred tank reactors (42). The oxygen mass transfer coefficient calculation associated with the solvent plant’s mass balance is presented in Table S1 in the Supplementary Information (available with the online version of this article), falling at the lower end of the range of mass transfer coefficients reported by Peterson et al. (42). Details of the experimental gas fermentation data is presented in Table III; a more detailed explanation of the experimental procedure can be found in Bommareddy et al. (28).

Significant heat integration makes the solvent plant self-sufficient in electricity and both low and medium pressure steam. Furthermore, process water recovered from distillation and the steam condensate is recycled to reduce the water make-up burden.

The process flow diagram for conventional renewable electricity generation, used to value the black liquor, is presented in Figure 3. An additional steam turbine is required to produce the renewable electricity for sale, relying upon the existing multi-effect evaporators, air compression and Tomlinson boiler. Superheated steam at 9000 KPa and 480ºC is used in the steam turbine (44). The medium pressure steam exiting the turbine is used in the multi-effect evaporators to concentrate the excess black liquor to 75% and to preheat the auxiliary air supplied to the Tomlinson boiler. Similarly, the associated electricity demand for the air compressor and pump is provided by the electricity generated. Resultantly, through conventional renewable electricity generation, the excess black liquor produces 138 GWh year^–1 for sale to the grid.

Fig. 3

The mass and energy balance associated with the rigorous process simulation informs the capital cost, fixed operating cost and variable operating cost estimation. For the capital cost estimation, major equipment purchase costs were estimated using the models from Seider et al. (45), with the exception of the turbo-expander (46). Three different methods are used to calculate the FCI, owed to differences in the estimation methods. These three methods are designated as: the NREL method outlined in the 2011 NREL report (27), the Towler and Sinnott (TS) method taken from Chemical Engineering Design (47) and the Hand method detailed in Sustainable Design Through Process Integration (48). The calculation basis of the three methods is presented in Table IV.

For all three methods, the calculated equipment purchase costs are multiplied by an installation factor to obtain the inside battery limit (ISBL) installed costs. Both the NREL and Hand methods use installation factors dependant on the equipment type, whereas the TS method uses a universal multiplier. All installed equipment costs were adjusted to 2019 costs using the Chemical Engineering Plant Cost Index of 607.5 (49). A location factor of 0.51 was used for China (using indigenous materials), based on the 2003 location factor of 0.61 (47), updated to 2019 via the Chinese Yuan to US dollar exchange rate.

Three methods were used to calculate the fixed operating costs as summarised in Table V. As before, the NREL method (27) and the TS method (47) were employed. However, as the Hand method is solely for FCI, the third was the taken from Coulson and Richardson Volume 6 (50). Variable operating costs were estimated based on the costs detailed in Table VI, subject to annual inflation as outlined in Table VII.

Time series analysis was used to forecast the long-term average price of isopropanol and acetone. Takens’ theorem was used as the basis for this analysis (53). Takens’ theorem states that for a deterministic system, the underlying state variables that created the time series are embedded within the data. Using this theorem a deterministic, dynamic system can be reconstructed based on the observed time series. Forecast models constructed using the embedded state variables assume that the market drivers underpinning the trajectory of the state variables in phase space remain largely unchanged. An embedding dimension of 10 was used to reconstruct the isopropanol and acetone price models from monthly average price data obtained from the Intratec database (54). In this work, a radial basis function neural network (RBFNN) containing eight neurons was used as a model to predict the future commodity prices. The network was trained as a one step ahead predictor by minimising the mean square error of the difference between the actual and predicted prices. Once trained, the network was evaluated (tested) in free run mode, where successive predicted prices (outputs) become inputs to the RBFNN. The confidence limits corresponding to the trained RBFNN were calculated as a reliability measure of the prediction as per the work undertaken by Leonard, Kramer and Ungar (55). The benefit of using an RBFNN is that the resultant forecast price is an impartial product of the dataset’s underlying state variables.

The long-term average price for renewable electricity sales was taken as US$0.109 kWh^–1 as per the biomass subsidy in China (56). This is used to inform the renewable electricity project to value the black liquor and for the excess electricity generated by the solvent plant.

The cost models from Section 2.3 and the product price forecast models from Section 2.4 inform the investment analyses. The black liquor is costed at its utility value, calculated as the foregone NPV from generating renewable electricity. Resultantly, the NPV for the solvent plant is calculated by subtracting the NPV of renewable electricity generation. The investment analysis parameters used are detailed in Table VII.

A sensitivity analysis was conducted using a Monte Carlo simulation based on the cost parameters in Table VIII, creating an uncertainty framework. The cost parameters were taken from (47), with the exception of renewable electricity sale price where the upper limit for the long-term average price was capped at the current biomass subsidy in China, US$0.109 kWh^–1. This limit was applied due to the decreasing trend in renewable electricity subsidies (58). In contrast, the long-term average prices for isopropanol and acetone were varied ±30% from the forecast price. This provides a stochastic counter to the assumption used to determine the forecast prices: that the deterministic market drivers underpinning the trajectory of the state variables remain largely unchanged. However, given that market drivers are subject to change, the long-term average price may be banded with an equal likelihood of being higher or lower than the forecast price.

A uniform distribution for these parameters was used and varied for the solvent plant and conventional renewable electricity generation (used to value the black liquor). All the cost parameters in Table VIII, other than labour and electricity, were varied independently. 2000 simulations were run, stochastically varying the parameters within the defined lower and upper limits to produce a probability distribution of the solvent plant’s NPV.

A cradle-to-gate LCA model was developed using the ecoinvent 3.6 inventory database, following ISO Standards 14040 (59) and 14044 (60). GHG emissions were calculated based on the most recent Integrated Pollution Prevention and Control 100-year global warming potential (GWP) factors to quantify GHG emissions in terms of carbon dioxide equivalents (CO_2eq) (61). Functional units were defined as 1 kg isopropanol, 1 kg acetone and 1 kWh of electricity. In line with the investment analysis, the LCA model considers the net electricity output of solvent plant by subtracting the foregone electricity from combustion of black liquor at the pulp mill. Life cycle environmental impacts are allocated between these three products using both economic and energy allocation. The GHG emission rate for the external process inputs: cooling water, process water and ammonia were taken from the ecoinvent 3.6 inventory database using the allocation at the point of substitution system model (62), whereas electricity was taken as the 2018 China electricity mix (63). The bio-based solvents isopropanol and acetone sequester biogenic carbon dioxide and hence are credited with a negative GHG emission based on their carbon content. Downstream activities, including the use and end-of-life of isopropanol and acetone products are not considered. These activities are assumed to be identical to those of conventional isopropanol and acetone, given that they are chemically and functionally identical, and therefore have no influence on the relative GHG emissions of renewable and conventional solvent products.

The solvent plant (Figure 2) produces three products, summarised in Table IX. The contribution of each product to the plant’s income is also presented. Whilst isopropanol contributes to almost half the solvent plant income the renewable electricity fraction is the second highest contributor, highlighting the significant amount of renewable electricity generated by the solvent plant.

The investment analyses for the solvent plant and conventional renewable electricity generation are detailed in Tables S9 and S12, as per the investment analysis parameters presented in Table VII. The NPV for conventional renewable electricity generation represents the utility value of the black liquor, valued at US$73 million (Table S12). This is subtracted from the NPV of the solvent plant (US$115 million) to produce the cumulative NPV presented in Figure 9. For the nominal TEA model inputs, the solvent plant’s net cumulative NPV is US$42 million.

Fig. 9

Given the conceptual stage of the TEA, a Monte Carlo simulation was undertaken as per the uncertainty framework outlined in Table VIII. The produced probability distribution in Figure 10 avoids making an investment decision based solely on nominal TEA inputs. The cumulative probability curve presents a 70% probability that the solvent plant will achieve a net cumulative NPV between US$35 million and US$85 million, noting that no negative outcomes are predicted.

Fig. 10

Figure 11 summarises the outcome of the cradle-to-gate LCA for the solvent plant, compared to the conventional fossil derived processes; using both economic and energy allocation for the isopropanol, acetone and renewable electricity products.

Fig. 11

Both solvents achieve negative GHG emissions when produced via the solvent plant using economic and energy allocation. The GHG emission for the two allocation methods are comparable, indicating the price per unit energy (US$ MJ^–1) is similar for all three products. The negative emissions are an intrinsic outcome of the cradle-to-gate framework, which excludes the end use for the products. As the total GHG emissions of the solvent plant are lower than the overall biogenic carbon sequestered, negative GHG emissions are achieved for the solvent products.

The negative GHG emissions compare favourably to the conventional isopropanol (hydration of propene) and acetone (oxidation of cumene) processes. Additionally, the GHG emissions associated with the excess renewable electricity from the solvent plant also compare favourably to the electricity mix in China 2018). Furthermore, as the end use for the solvents remains the same regardless of the production method, the relative GHG emissions are valid beyond the cradle-to-gate framework.

As highlighted in the Introduction, the commercial implementation of gas fermentation is largely dominated by anaerobic fermentation. Therefore, it is important to compare the results to a best-in-class technology. In addition to successfully commercialising ethanol production via gas fermentation, LanzaTech have also investigated gas fermentation to produce acetone, a precursor to isopropanol (64). As such, LanzaTech’s investigation undertaken for the US Department of Energy (US DOE), in collaboration with Oak Ridge National Laboratory, USA, is used as a benchmark anaerobic process (65).

As highlighted previously, the primary differences between anaerobic and aerobic fermentation technologies are inherent to the C1 fixation metabolic pathways. Strictly respiratory (aerobic) cell factories require air to be continuously fed into the bioreactor to facilitate carbon fixation. In addition, owed to the intrinsic thermodynamic inefficiency of the Calvin-Benson-Bassham cycle employed by aerobic bacteria, an excess of low temperature heat is produced. As such, a conventional process flowsheet for aerobic fermentation employs operationally costly compressors and chillers. In contrast, for anaerobic fermentation there is a reduced chiller requirement and the compressor duty is less pronounced. Moreover, owed to the presence of oxygen, aerobic fermentations require the use of more costly stainless steel reactors and more complex process control systems. Whilst the latter is an intrinsic requirement of aerobic fermentations, in this work we have reconciled the increased utility demand of aerobic fermentation through process integration (28). This integration employs a heat pump to utilise the low temperature heat generated by aerobic fermentation to heat the SCWG reactor feed, removing the cooling water burden required by the bioreactors. Additionally, the compressor duty is fully supplied through the electricity generated upon letting down the SCWG reactor’s high-pressure gas product. As a result, the economic and LCA outcomes for the solvent plant should be comparable to anaerobic fermentation technology.

LanzaTech’s anaerobic study achieved a combined selectivity of 94.7% for ethanol and acetone, of which 57.3% was acetone (65). LanzaTech disclosed that by selling acetone at market prices they are able to sell coproduced ethanol at or below the US DOE’s 2022 target of US$3 GGE^–1 (GGE = gallon of gasoline equivalent) (66). Therefore, in this study, the price per GGE was calculated for the solvent products as a biofuel mix, with renewable electricity sold at the current market price. A value of US$2.87 GGE^–1 (Figure 12) was obtained, below the US DOE’s target, highlighting the competitiveness of the heat integrated aerobic solvent plant. Notably, neither isopropanol nor acetone are typically used for their fuel value, highlighted by their higher market prices. As such, the solvent plant is profitable as either a biofuel or commodity chemical facility.

Fig. 12

For LanzaTech’s anaerobic process, the cradle-to-gate LCA using energy allocation produced a calculated GHG emission of –1.9 kg_CO2eq kg^–1_{acetone + ethanol} for a heat integrated scenario (see Table S13 for calculation). In Figure 12, the LCA for the solvent plant is presented, indicating a net GHG emission of –2.04 kg_CO2eq kg^–1_{isopropanol + acetone}, which is in line with LanzaTech’s study (Figure 12). Resultantly, from both the TEA and LCA results, the greater thermodynamic efficiency of the anaerobic Wood-Ljungdahl C1 fixation pathway over the aerobic Calvin-Benson-Bassham Cycle is not evident for the heat integrated solvent plant.

At present, primary iron metal is commonly produced through the CO₂ intensive carbothermic reduction of iron oxides in blast furnaces at a temperature of around 1600°C. Since carbon is used as both reducing agent and fuel for the process, blast furnace pig-iron cannot eliminate its CO₂ emissions. Therefore, in recent years, carbon-free electrochemical processes have been widely investigated as potential green alternative routes for the production of iron and iron-base alloys (1–4) (assuming renewable electricity).

A large project that has been focused on alternative ways of producing iron was Ultra Low CO₂ in Steelmaking (ULCOS) in which new smelting reduction concepts were studied. The ULCOWIN electrolytic production of iron from suspensions of iron oxide particles in a highly concentrated sodium hydroxide solution at 110°C was demonstrated at laboratory scale. It has been shown that the iron particles are reduced in the solid state, which differs from the conventional electrowinning processes where the metal is deposited through the reduction of dissolved metal cations. Previous works with this process achieved high Faradaic yield (80–95%) (1–4).

Based on the above-mentioned studies, the technology for alkaline pulp iron electrowinning is being studied for the first time from a secondary mineral source, namely bauxite residue from the alumina refining industry. This technique is referred to in the SIDERWIN project (5), which aims among others to produce iron from alternative low-grade iron sources, currently incompatible with the conventional steel making processes.

Bauxite ore is treated within the Bayer process to produce metallurgical grade alumina which is the raw material for aluminium production. Bauxite ore depending on its origin contains 40–60% alumina and the rest is a mixture of iron (20–30%), silicon and titanium oxides. When bauxite ore is treated with caustic soda, the aluminium hydroxides or oxides contained within are solubilised, with approximately 50% of the bauxite mass being transferred to the liquid phase, while the remaining solid fraction constitutes the bauxite residue, often termed as ‘red mud’ due to its colour. Depending on the grade of the bauxite ore used, bauxite residue, on a dry basis, is produced from 0.9 to 1.5 mass ratio to the alumina product (6). The high volume and alkalinity of this byproduct make its valorisation a major challenge worldwide (7).

Bauxite residue is an untapped secondary raw material source considering the presence of valuable substances such as iron (30–45 wt%), aluminium (15–25 wt%), silicon, calcium, titanium and sodium oxides as well as smaller concentrations of critical or industrially important elements such as rare earth elements (REEs) (mainly cerium, lanthanum, scandium, yttrium and neodymium), vanadium, chromium and others (6). The recovery of the major metals from bauxite residue has never yet been implemented while a lot of processes have been proposed. Despite the laboratory-scale success of much of the work so far, currently the industrial utilisation of bauxite residue is estimated at just 2–4 million tonnes, accounting for less than 3% of the annual bauxite residue production (8).

The main constituent of bauxite residue is iron oxide and it can make up to 45% of the mass of the bauxite residue. In fact, the red colour of bauxite residue is caused by iron(III) oxides (mostly haematite, Fe₂O₃) (9). In general, due to its high alkaline (Na₂O ≈ 10%) and its titanium (TiO₂ ≈ 4–15%) content, bauxite residue is not suitable for use as an iron ore substitute in blast furnaces.

Bauxite residue reductive smelting processes can be applied by several technologies (Corex, Finex^®, HIsmelt, Romelt, AusIron and electric arc furnace (EAF)) for the production of pig iron (10, 11). So far, two methods have been applied at pilot scale for bauxite residue reductive smelting: the Romelt method (12) and the EAF (13–16). The Moscow Institute of Steel and Alloys (MISA) (Russia), with National Aluminium Company Ltd (NALCO) and the joint venture company Romelt-Steel Authority of India Ltd (RSIL) (India), studied the pyrometallurgical process of bauxite residues using the Romelt method (12). The advantage of this method is that materials can be used with moisture levels of up to 10 wt%. The main disadvantage is the high energy consumption and the poor quality of pig iron with a high concentration of sulfur and phosphorus (10). In EAF reductive smelting, a mixture of bauxite residue, carbon and fluxes is treated at 1500–1700°C to form pig iron with higher than 95% iron recovery (6, 13, 17). Recovery of residual iron can be further improved by later magnetic separation in slag dust (18). Post-melting slag can be used to produce rockwool or building materials (6, 19) and for the recovery of non-ferrous metals and REE (9, 18, 20, 21). However, such methods have not been industrialised as they are not competitive to established iron and steel making processes.

The present paper describes a highly promising electrochemical method for sustainably extracting the iron from bauxite residue, in an alkaline environment, which is in-tune with the alkaline environment of the Bayer process.

The process is conducted in an electrolysis cell consisting of a borosilicate glass beaker (250 ml) closed with a specially configured cylindrical silicon bung (45 cm diameter). The experimental apparatus is shown in Figure 1.

Fig. 1.

A three-electrode configuration was used; the cathode was a rectangular shaped stainless steel V2A plate (110 mm height, 10 mm length, 1 mm width), the anodes were two rectangular shaped nickel plates (100 mm height, 10 mm length, 2 mm width). All electrodes were centred according to the silicon bung’s cylindrical axis in specially configured holes so that the electrodes were in certain positions and at fixed distance to each other. The working electrode’s surface area that was immersed in the solution was defined to be 8 cm². The reference electrode that was used was a commercial Hg | HgO | NaOH (1 M) electrode (RE-61AP, ALS Co Ltd, Japan) which was immersed in a distinct glass.

Bauxite residue was supplied by MYTILINEOS-Aluminium of Greece. Samples were solubilised via fusion method according to which a quantity of bauxite residue remained at 1000°C for 1 h with a mixture of Li₂B₄O₇/KNO₃ followed by direct dissolution in 6.5% nitric acid solution. Chemical analysis was performed by atomic absorption spectroscopy (AAS) with the use of a PerkinElmer 2100 atomic absorption spectrometer. The chemical analysis of the samples used in this study is shown in Table I. The chemical analysis showed that Fe₂O₃ is the main content of bauxite residue.

Mineralogical analysis of bauxite residue (Figure 2) showed that the main iron oxide phase is haematite while a small amount of goethite also exists. The haematite to goethite mass ratio in bauxite residue is 4.2 (22).

Fig. 2.

Prior to each experiment, the stainless-steel cathode and the nickel anodes were polished with sandpaper and were rinsed with demineralised water. Furthermore, the cathode was weighed. The electrolyte was a 50 wt% NaOH aqueous solution corresponding to molarity of 25 mol kg^–1, to which was added 10 wt% bauxite residue solid particles. Typically, a mixture of 152.4 g solid NaOH with purity higher than 99% (CHEM-LAB NV, Belgium) and 152.4 g demineralised water were mixed under stirring for about 30 min. After the electrolyte’s homogenisation, 33.9 g of bauxite residue were slowly added to the solution within 5 min.

The slurry temperature was measured via a probe that was wrapped in a polytetrafluoroethylene (PTFE) shrink tubing to avoid current leakage. The slurry was stirred using a 1 cm ringed cylindrical magnetic bar and magnetic hot plate (IKA^® RCT basic, IKA Works Inc, USA) at a rotational speed of 500 rpm to keep bauxite residue particles suspended. The small size of the magnetic bar was chosen to minimise the attraction of magnetite particles to the stirrer. The cathode, the anodes and the reference electrode were connected via their respective terminal to a potentiostat (SourceMeter^® 2461 Series, Keithley, Tektronix Inc, USA).

Cyclic voltammetry and electrolysis tests, under chronopotentiometry mode, were performed in a three electrode cell connected to a potentiostat (2461 Series, Keithley) and the obtained experimental data were analysed. Regarding cyclic voltammetry, it should be noted that both working and counter electrode were platinum wires while the reference electrode was the above mentioned Hg/HgO commercial electrode.

Electrolysis tests were performed under galvanostatic mode. The duration of the experiment was 2 h. After the end of each experiment, the cathode was thoroughly rinsed with distilled water in order to remove the remaining electrolyte, and was dried at 100°C for 24 h, before being weighed. The difference between the mass of the cathode prior to and after the end of electrolysis was considered as the mass of the deposit. This value was used to deduce the current efficiency according to Faradaic law.

2.1 Determination of Metallic Iron

3.1 Cyclic Voltammetry

3.2 Effect of Current density

3.3 Effect of Temperature

The present work has demonstrated the possibility to electrochemically reduce iron from bauxite residue in alkaline pulps. The process temperature proved to be the most crucial parameter that substantially affects the Faradaic process efficiency. At the current level of process development, a Faradaic efficiency of 71.58% was achieved at pulp density of 10 wt% bauxite residue in a 50 wt% NaOH solution at 130°C. The applied current has to create a cathodic potential higher than –1.4 V vs. Hg/HgO (in 1 M NaOH) in order to avoid the hydrogen evolution which takes place at cathodic potentials lower than –1.4 V.

The cathode in the case of bauxite residue pulps electrolysis is not uniformly covered by electroactive iron oxide particles (mainly haematite) and therefore there is not a uniform cathodic potential on the whole cathode surface. This affects the current efficiency of the process which is always substantially lower than that observed in pure haematite ore electrolysis.

Bauxite residue is produced as a byproduct of the Bayer process, an alkaline leaching process taking place at temperatures 120–250°C. Therefore, the present process has great potential for integration in the established Bayer process as a symbiotic step to valorise the (currently wasted) iron portion of the bauxite ore.

The adsorption of air and water contaminants on an activated carbon surface is frequently used in air and water treatment systems. Conventional biological treatment processes are efficient in removing biodegradable organics. These processes, however, have limited ability to remove lignins, humic substances, pesticides and residual colour and odour-producing organic compounds. The performance of activated carbon in removing such organic compounds has been proven. Activated carbon filters used for water treatment in homes are usually made of either powdered activated carbon (PAC) or granular activated carbon (GAC). PAC is made of powdered carbon with a particle size of <0.18 mm and diameter in the range of 0.25–1.15 mm. It is generally used in batch type reactors and is subsequently filtered off. It is used in the treatment of liquids and cleaning of flue gases. GAC is produced from 0.2–5 mm particle size irregular carbon. For liquid and gas phase substances, it is applied on adsorption columns (1–3).

The increase in production and accompanying energy demand which has emerged in the last century with the increasing population is becoming important. The environmental impacts caused by the increase in production draw attention and higher consumption creates problems for our limited natural resources. The necessity of treatment for the prevention of soil, air and water pollution is among the topics that are especially emphasised at international conferences. It is necessary for humanity to take joint decisions to increase the sensitivity of countries on environmental issues and to implement the necessary legal measures (4).

Biochar is a byproduct with high carbon content produced through the conversion of biomass into syngas by advanced thermal gasification using partial oxidation. Intensified gasification of biomass through partial oxidation is a recently developed, environmentally clean and renewable energy technology (5). Biochar obtained as a byproduct from gasification of woody biomass mainly contains carbon (85 wt%), nitrogen, hydrogen, phosphorus, calcium, magnesium and iron. Production of activated carbon from biochar is possible. To convert biochar into activated carbon, it is necessary to remove the specific contaminants to complete the activation.

Biomass is a renewable energy source but can cause environmental issues when it is not properly utilised. Therefore, it needs to be dealt using appropriate technology as compared to other renewable energy sources. Since thermal utilisation of biomass is carbon neutral, it does not cause global warming, and due to the quantities available it has high energy potential. Thermochemical methods are generally used to obtain energy from biomass. Known thermochemical methods are combustion, pyrolysis and gasification. The gasification process has several advantages over other thermal processes: the ability to dissolve material at lower reactor volumes, the formation of low amounts of contaminants and more efficient utilisation of the produced syngas. Compared to pyrolysis, it has the advantage of working autothermally without the need for external energy. Compared to other processes, gasification of woody biomass is stated to be one of the most suitable options for optimising the conversion efficiency of the fuel’s chemical energy (6–7). The stoichiometric reaction of woody biomass with partial oxygen results in flammable product gases and thus gasification. The basic equation of the gasification of biomass is as in Equation (i):

(i)

Various types of reactors are available for gasification. To not digress from the subject, the kinetics of gasification reaction was not elaborated. Fixed-bed biomass gasification reactors are used for syngas generation and biochar production. Fixed bed gasification reactors efficiently convert biomass to syngas and are the type of reactor suitable for the production of biochar as a byproduct with high carbon content (85–90 wt% carbon) (8–9).

The biomass fed into the gasification system passes through four sections in the reactor and is converted into syngas and biochar as a byproduct. The biomass passes through the drying, pyrolysis, reduction and oxidation zones in the reactor and is converted into biochar at elevated temperatures (800–1100ºC). The amount of biochar produced depends on the biomass type and the operational conditions. 10–20% of the biomass fed to the reactor is produced as biochar (10–12).

Recently, researchers have begun to be interested in biochar produced as a byproduct from biomass by thermal conversion pathways (>700ºC), such as pyrolysis, thermal carbonisation, flash carbonisation and gasification (13). Many studies have suggested that biochar can be used in CO₂ adsorption, soil remediation and air pollution removal (14–16). Additionally, the use of the syngas produced by thermochemical conversion of biomass as a renewable energy source provides an advantage while the efficiency potential of biochar produced as a byproduct will make these systems very attractive (17).

Physical activation improves the surface pores of biochar and affects the chemical properties of the surface (such as surface functional groups, hydrophobicity and polarity). The physical activation methods used for biochar are generally steam and gas activation.

The steam activation of the biochar is usually carried out after the thermal carbonisation of the biomass. The surface porosity of biochar increases after pyrolysis. Also, with steam activation, the activated biochar will gain higher porosity. The chemical formulae are shown in Equations (ii) and (iii) (18):

(ii)

(iii)

With the H₂O (steam) and carbon reaction during the activation process, three outcomes emerge: (a) removal of the volatiles and tar from the surface; (b) formation of new micropores; and (c) expansion of existing pores (19–20).

Studies on steam activation have shown that it increases the biochar’s surface area and micropores. For example, the surface area of biochar (136–793 m² g⁻¹) obtained by rapid pyrolysis (800ºC, 45 min) was increased as a result of steam activation. All these measurements were performed with BET isotherm (21). In another study, biochar (700ºC) from tea residue biomass was activated by steam and its surface area was obtained as 576.1 m² g⁻¹, pore volume as 0.109 cm³ g⁻¹ and pore diameter as 1.998 nm (before activation, the surface area was 342.2 m² g⁻¹, pore volume 0.022 cm³ g⁻¹ and pore diameter 1.756 nm) (22). The steam activated biochar formed at 700ºC indicated the highest absorption capacity (37.7 mg g⁻¹) at pH 3, with a 55% growth in absorption ability compared to non-activated biochar produced under the same conditions. Consequently, activation with steam has potential to enhance the adsorption capability of biochar. At 700ºC, produced biochar derived from plant-based biomass detected relatively low surface areas (9.27 m² g⁻¹). Due to the formation of tar during thermal decomposition, the pores in the biochar are blocked (23). Steam activated biochar derived from bamboo waste had a larger surface area compared to non-activated bamboo biochar. Optimum conditions for activation were 850ºC and 120 min activation time. Under these conditions, the BET surface area of activated biochar was 1210 m² g⁻¹ and total pore volume was 0.542 cm³ g⁻¹. This study showed that bamboo waste could be used to arrange new micropores in activated biochar through steam activation (24). In further research, activation of rice husk biochar was carried out at 800ºC using steam. Micro- and mesoporous structured biochar were produced and 1365 m² g⁻¹ surface area was obtained at the end of 15 min (25).

Another paper presents a study into the effect of different activation conditions and adsorption characteristics of biochar evaluated from tyre rubber waste. Steam was used as an oxidising agent and total micropore volumes and BET surface areas increased to 0.554 cm³ g⁻¹ and 1070 m² g⁻¹, respectively. Consequently, steam was observed to generate a narrower extensive microporosity (26). An experimental test represents the production of activated biochar from barley straw using steam activation. Activation was conducted to maximise the micropore volumes and BET surface area of the biochar. Optimal conditions for steam activation were a hold time of 1 h at 700ºC. The micropore volume and surface area of the activated biochar were 0.2304 cm³ g⁻¹ and 552 m² g⁻¹, respectively (27). A further study investigated activated biochar produced from date stone wastes by steam activation. The effect of activation hold time on surface textural structure properties of raw date stone and biochar were studied. The results indicated the presence of cellulose and hemicellulose in the raw material, and the predominance of carbon content. The highest microporous volume was 0.716 cm³ g⁻¹ and specific surface area was 635 m² g⁻¹, obtained through biochar activated under steam at 700°C for 6 h (28). Another study was conducted to investigate the effect of sulfur in activated biochar prepared from apricot stones by steam activation. The activation temperature and time tested were in the ranges of 650–850ºC for 1–4 h. The experimental results revealed that the surface area of the biochar was 1092 m² g⁻¹ at activation conditions of 800°C for 4 h. The experimental results indicated that commercial production of porous activated biochar from apricot stones is reasonable (29). The significance of the nature and composition of biomass was demonstrated by the steam activation of three different biomass sources including wheat straw, coconut shell and willow (30). For these three biomass substances, activated carbons with specific surface areas of respectively 246 m² g⁻¹, 626 m² g⁻¹ and 840 m² g⁻¹ were produced.

Although physical activation using steam significantly increases surface area and porosity, there is a disadvantage of using steam activation. This is the loss of aromaticity and polarity in steam activated carbons compared to genuine biochar (31). However, using a combination of CO₂ and steam activation has been reported to produce activated carbons with better surface area and pore structure compared to using only CO₂ or steam for activation. In one study, activated carbon was produced from biochar obtained from olive kernel biomass using CO₂, steam and a combination of CO₂ and steam. A combined CO₂ and steam activation under similar experimental conditions was reported to produce a higher surface area (1187 m² g⁻¹) compared to the specific surface area obtained by using only CO₂ (572 m² g⁻¹) or steam (1074 m² g⁻¹) (32). From the aforementioned studies, it can be concluded that the reaction of steam with carbon occurs in a shorter period compared to the reaction of CO₂ with carbon. While generally micropore activated carbons are produced by using CO₂ in physical activation, meso- and macropores are formed in the structure during activation with steam (33). The reason for the difference in pores is attributed to the conversion of developed micropores into wide meso- or macropores and the faster reaction of the fixed carbon in the biomass structure of the steam. Also, it is possible to produce more pores using steam as it can penetrate the inner surface of the fixed carbon. On the other hand, CO₂ stagnates in its reaction with the fixed carbon and therefore more homogeneous micropores are added to the structure due to activation using CO₂. Determining the specified exposure time of the biomass at high temperatures with the activation agent is a critical decision to achieve high surface properties. The significance of biochar steam activation time was emphasised in some studies. It has been determined that there is a decrease in the surface pores of biochar exposed to long activation time (34).

Chemical and physical processes are needed to increase the utilisation value of biochar. Several processes are required such as separation of biochar into suitable granule size, obtaining carbon black, activation of pores by chemical processes and the sizing of activated carbon. The failure to use biochar produced as a byproduct of biomass gasification constitutes a disadvantage in all respects considering the spread of such technologies. Biochar produced as a byproduct by gasification of woody biomass is only used as a soil conditioner in small amounts. It is generally burned inefficiently in combustion boilers. To this end, this study investigates the conversion of biochar into activated carbon, which is a widely used valuable product. Several studies have emphasised that surface area and pore volume increase with the conversion of biomass at high temperatures by thermal methods (35). Thus, it is envisaged in the present study that the process of converting biochar obtained by gasification at high temperatures (800–1000ºC) into activated carbon by specific activation methods may have many advantages. The need for extra energy from external sources such as utilisation of fossil fuels to reach these temperatures in conventional carbon conversion systems causes high operational costs. Considering all this, it is expected that high quality activated carbons obtained from biochar will be available in the market at low costs.

In this experimental study, oak woodchip was gasified in an updraft gasifier. Byproduct biochar was activated in the activation unit integrated into the gasification system.

The proximate and ultimate analyses of the oak woodchips were conducted, and the results are presented in the Results and Discussion section. Proximate and elemental analyses were performed to determine the usability of biochar produced after the gasification of biomass for the production of activated biochar. Proximate analyses of moisture (ASTM D7582-12), ash (ASTM E1755-01(2020)), fixed carbon (ASTM D3172-13) and volatile matter (ASTM D7582-12) were made according to standard methods. The lower heating value was determined using ASTM D5865-13 standard method. The amounts of carbon, hydrogen, nitrogen and sulfur were determined in the element analyser and the amount of oxygen was determined using the standard ASTM D3176-09.

Using a thermal analyser, the thermogravimetric carbonisation analysis of oak woodchips was performed. Approximately 10 mg samples with an average particle size of 0.25 mm were heated at 10ºC min⁻¹ from room temperature to 800ºC under nitrogen flow. Throughout the measurements, the nitrogen flow was kept constant at 10 cm³ min⁻¹.

Activation of the biochar produced in the gasification system in the Gebze Technical University (GTU) Gasification Laboratory was performed. For the preparation of the biochar and activated biochar, oak woodchips were first carbonised by gasification at 800–1000ºC for 1–2 h. Then, the resultant biochar produced as a byproduct from the gasifier was activated by steam at three different temperatures (700ºC, 750ºC and 850ºC) utilising thermal heat generated from a syngas burner at different activation times (30 min, 60 min and 90 min). 5.5 kg byproduct biochar was used in each run. During the final process, the reactor was cooled under inert injection of nitrogen gas and then the activated carbon was removed from the reactor and weighed in order to determine the burn-off undergone in the reaction. For steam activation, a stainless-steel vertical reactor illustrated in Figure 1 was used and integrated to the system to heat each 5.5 kg biochar sample. Throughout the experiment, an exact heating rate and steam flow rate (~20ºC min⁻¹, 1.3 kg min⁻¹, respectively) were applied. Referring to Equation (ii), the optimal stoichiometric steam amount was determined and calculated for the system.

Fig. 1

The nitrogen adsorption or desorption isotherms were determined at 77 K by means of an automatic adsorption instrument to identify the textural properties of the produced biochar and activated biochar. Before the gas adsorption measurements, the samples were degassed at 300ºC under vacuum for 5 h. The N₂ adsorption isotherm was achieved over a relative pressure, P:P₀, ranging from roughly 10⁻⁶ to 1. The BET and t-plot methods were employed to determine the surface area, micropore surface area and pore volume of the biochar and activated biochar, respectively. Relative pressures in the 0.01–0.15 range were applied to evaluate the BET surface areas. The total pore volumes (V_t, cm³ g⁻¹) were considered to be the liquid volumes of N₂ at high relative pressure near unity (~0.99) (36–37).

SEM analysis was performed to investigate the surface, textural, porosity and structural properties of activated biochar produced under different conditions. SEM analyses were taken by enlarging ×500 to observe changes in surface morphological structure before and after activation.

To examine the crystal structure, XRD profiles of each sample were obtained at room temperature, using a copper Kα X-ray source, under 40 kV and 30 mA analysis conditions. Diffraction data were taken on a scale ranging from 2θ = 0–90º.

In order to qualitatively determine surface functional groups, FTIR spectra were obtained at room temperature, with the support of diamond orbital attenuated total reflection (ATR) accessory, by scanning 128 times in the range of 500–4000 cm⁻¹ band at 4 cm⁻¹ separation sensitivity. Samples were placed directly on the diamond crystal and were analysed by applying pressure to allow it to fully interact with the diamond crystal.