Microbial synthesis gas utilization and ways to resolve kinetic and mass-transfer limitations

Accepted Manuscript Review Microbial synthesis gas utilization and ways to resolve kinetic and mass-transfer limitations Muhammad Yasin, Yeseul Jeong, Shinyoung Park, Jiyeong Jeong, Eun Yeol Lee, Robert W. Lovitt, Byung Hong Kim, Jinwon Lee, In Seop Chang PII: DOI: Reference:

S0960-8524(14)01624-1 http://dx.doi.org/10.1016/j.biortech.2014.11.022 BITE 14228

To appear in:

Bioresource Technology

Received Date: Revised Date: Accepted Date:

23 September 2014 6 November 2014 8 November 2014

Please cite this article as: Yasin, M., Jeong, Y., Park, S., Jeong, J., Lee, E.Y., Lovitt, R.W., Kim, B.H., Lee, J., Chang, I.S., Microbial synthesis gas utilization and ways to resolve kinetic and mass-transfer limitations, Bioresource Technology (2014), doi: http://dx.doi.org/10.1016/j.biortech.2014.11.022

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Microbial synthesis gas utilization and ways to resolve kinetic and mass-

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transfer limitations

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Muhammad Yasina, Yeseul Jeonga, Shinyoung Parka, Jiyeong Jeonga, Eun Yeol Leeb,

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Robert W. Lovittc, Byung Hong Kimd, Jinwon Leee, In Seop Changa,*

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a. School of Environmental Science and Engineering, Gwangju Institute of Science and Technology, Republic of Korea

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b. Department of Chemical Engineering, Kyung Hee University, Gyeonggi-do 446-701, Republic of Korea

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c. College of Engineering, Swansea University, Swansea SA2 8PP, United Kingdom

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d. Fuel Cell Institute, National University of Malaysia, 43600 UKM, Bangi, Malaysia

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e. Department of Chemical and Biomolecular Engineering, Sogang University, Seoul, 121-742, Republic of Korea

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*Corresponding author: [email protected]

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Abstract Microbial conversion of syngas to energy-dense biofuels and valuable chemicals is a

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potential technology for the efficient utilization of fossils (e.g., coal) and renewable resources

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(e.g., lignocellulosic biomass) in an environmentally friendly manner. However, gas-liquid mass

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transfer and kinetic limitations are still major constraints that limit the widespread adoption and

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successful commercialization of the technology. This review paper provides rationales for syngas

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bioconversion and summarizes the reaction limited conditions along with the possible strategies

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to overcome these challenges. Mass transfer and economic performances of various reactor

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configurations are compared, and an ideal case for optimum bioreactor operation is presented.

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Overall, the challenges with the bioprocessing steps are highlighted, and potential solutions are

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suggested. Future research directions are provided and a conceptual design for a membrane-

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based syngas biorefinery is proposed.

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Keywords: Syngas fermentation; C1 biorefinery; Acetogens; Mass transfer limitations; Hollow

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fiber membrane bioreactor

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1. Introduction A tremendous increase in energy demand has been witnessed across the globe due to rapid

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increases in the world population and growing industrialization. The primary energy

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requirements are fulfilled by utilizing petroleum reserves which are on the verge of extinction

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and are estimated to be exhausted in less than 50 years; if continuously used at the present

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consumption rate (Demirbas, 2007). Using fossil energy resources such as coal results in the

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enormous release of hazardous and toxic compounds that create an unhealthy environment for

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the biota on the earth. However, coal reserves, which are six-fold greater than petroleum

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reserves, should be considered as a short term energy source to be used in an environmentally

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friendly manner (Kim & Chang, 2009). Petroleum reserves are also able to be saved by

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introducing alternative biofuels from carbohydrates present in lignocellulosic biomass, while

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environmental concerns caused by coal can be addressed by developing effective and

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environmentally friendly processes. One possible alternative is to use a hybrid process that

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involves the simultaneous conversion of organic matter (from coal and biomass) into synthesis

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gas (syngas) via gasification followed by biological conversion utilizing microorganisms that are

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able to convert major components of syngas (CO and H2) into multi-carbon compounds (Henstra

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et al., 2007). Agricultural residues and woody biomass are not applicable for direct microbial

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conversion due to poor biodegradability of the lignin moiety called “recalcitrance”; however,

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lignin is one of the most energy rich components and creates 10 to 20% of the entire biomass

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present in nature (Daniell et al., 2012). Lignin content of biomass can be utilized effectively in

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syngas-based biorefinery.

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Syngas is used as a feed stock for the production of useful chemicals such as acetic and

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butyric acids and energy dense biofuels (i.e., ethanol and butanol) either by chemical catalytic

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conversion (e.g., Fischer-Tropsch (FT) Synthesis) or biological conversion routes. Biological

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syngas fermentation is considered to be more attractive due to its several inherent merits

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compared with the biochemical approach (enzymatic hydrolysis) and the FT process. Biological

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catalysts ferment syngas into liquid fuels effectively and efficiently compared with the use of

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chemical catalysts and require less energy and infrastructure set-up costs (Henstra et al., 2007).

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In addition, chemical composition differences of biomass are eliminated in the gasification

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process which enables the utilization of biomass from a variety of resources for biological

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conversion, which is typically called syngas fermentation. The microbial conversion of CO or H2

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to multi-carbon compounds occurs through the “Acetyl-CoA pathway”, which uses CO2 as a

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terminal electron acceptor for ATP (adenosine triphosphate) generation; therefore, microbial

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“respiration” is technically more accurate than “fermentation”. However, “syngas fermentation”

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is widely accepted and used by biotechnologists and process engineers (Yasin et al., 2014).

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Biological production of gaseous fuel (i.e., hydrogen) is also possible through water gas shift

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(WGS) reaction. Many thermophilic bacteria and archea have already been isolated that produce

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H2 through biological WGS reactions using CO as a reactant (Henstra et al., 2007). Due to high

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specific energy, the process development for biological conversion of CO into H2 has equal

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importance in achieving the long term goal of future CO biorefinery.

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The production of sustainable and renewable biofuels and valuable chemicals can be

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realized by using versatile feed stock resources for syngas fermentation; however, low

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achievable cell density in fermentation media due to the poor mass transfer rate of sparingly

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soluble gases (CO and H2) is still big hurdle in the successful design and operation of large scale

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process (Bredwell et al., 1999; Daniell et al., 2012). Chang et al. (2001) reported that low cell

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concentrations are associated with poor mass transfer of the gaseous substrates, while gas-liquid

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mass transfer (GL-MT) is the most dominant factor in fermentation broths containing higher cell

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concentrations (Abubackar et al., 2011). Thus, designing and operating commercial syngas

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fermentation facilities at high GL-MT conditions is strongly desired.

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Generally, GL-MT limits the conversion rates in bioprocesses that use sparingly soluble

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gases as key components such as; carbon and energy sources (CO or H2 in homoacetogens; CH4

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in methylotrophs) or electron acceptors for ATP generation (O2 in aerobic respiration). The

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utilization of these sparingly soluble substrates is often a mass transfer-limited process when

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dissolved gases are not sufficiently large to support microbial requirements that are expressed by

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two terms, specific substrate consumption rate (
 ) and cell concentration (X). Vega et al. (1989b)

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explained multiple steps during gas to liquid mass transfer phenomenon, which requires the

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absorption of a gaseous substrate across the gas-liquid interface (I), the transfer of the dissolved

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gas to the fermentation media (II) and diffusion through the culture media to the cell surface.

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Although the relative magnitudes of the mass transfer resistances depend on the composition and

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rheological properties of the liquid, mixing intensity, bubble size, interfacial adsorption and other

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factors, the most sparingly soluble gases utilized in the biochemical reaction cause major

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resistance in the liquid film around the gas-liquid interface (Klasson et al., 1992; Munasinghe &

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Khanal, 2010a; Vega et al., 1989b).

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To date, much effort has focused on increasing the volumetric GL-MT coefficient (kLa), in

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gas-utilizing bioprocesses. When gas is sparged through a liquid, the kLa principally depends on

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the size and number of bubbles present, which are affected by many factors such as agitation

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speed (Robinson & Wilke, 1973), gas flow rate (Yagi & Yoshida, 1975), reactor geometry

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(Klasson et al., 1992) and the nature of the liquid (Bredwell et al., 1997). Most of these methods

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rely on increasing the interfacial surface area available for the mass transfer and boost the bubble

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breakup by increasing the agitator’s power input to volume ratio. This approach, however, may

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not be practical on commercial scale due to high energy as well as infra cost requirements. To

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achieve energy efficient and high mass transferred conditions, alternative bioreactor

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configurations have been investigated for syngas fermentation (Bredwell et al., 1999;

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Munasinghe & Khanal, 2010b; Orgill et al., 2013). The efficiency of mass transfer rate in these

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reactors was observed through hydrodynamic conditions within reactors by predicting the kLa

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(Munasinghe & Khanal, 2010b).

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The suitability of a proposed reactor system for GL-MT applications is founded on its

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capability to correlate with the reaction kinetics. A bioreactor for these gaseous systems must

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operate in either of two regimes (Klasson et al., 1992). The first case includes the mass transfer

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limited, which occurs when sufficient cell mass is produced in the system to react with more

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solute, but the mass-transfer rate cannot keep pace; therefore, the liquid phase concentration goes

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to almost zero, and the cell concentration and rate of consumption are limited by the ability of

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that particular reactor to transfer substrate. However, in the other case, sufficient gaseous

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substrate is delivered to the reaction system, but the cell concentrations are not sufficient to allow

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an equal consumption rate making the conversion process kinetically limited (Ungerman &

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Heindel, 2007). Then the liquid phase concentration increases to reach the saturation point,

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promoting the condition of no driving force for mass transfer. In addition, if the saturation

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concentration shows a substrate inhibitory effect on microbes (substrate inhibition),
 is

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retarded in that particular bioreactor.

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The application of various reactor schemes to achieve high mass transfer is more

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practical and viable. Among the reactors, stirred tank reactors (STR) offer linear increase in kLa

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by increasing the agitation speed which results in a cubic increase in the power requirement

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(Orgill et al., 2013). Bubble column reactors (BCR)/gas spargers also produce higher kLa

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through increase in substrate intake flow rates. However, the power required for fluid (gas or

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liquid) circulation through pumps increases linearly with an increase in the flow rate, causing a

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linear increase in the cost (Kim et al., 2010a). The economic attractiveness of a biological

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process lies in the development of optimal bioreactor design that permits high gas-liquid mass

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transfer, high cell concentrations and high product concentrations in short residence times under

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no substrate limiting and inhibitory conditions.

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Recent studies regarding membrane bioreactors (MBR) for syngas fermentation have

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suggested that reactors installed hollow fiber membranes (HFM) as gas diffusers as a potential

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alternative to the widely commercially employed STR and the less common BCR and trickle bed

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reactors (TBR) (Munasinghe & Khanal, 2010b; Orgill et al., 2013; Shen et al., 2014b; Yasin et

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al., 2014). The performance of hollow fiber membrane bioreactors (HFMBR) in terms of mass

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transfer outcompete conventional reactors. Downstream processing steps in syngas fermentation

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also require research attention for the successful commercialization of the technology.

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Pervaporation is a membrane-based technology that offers an opportunity for the separation of

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volatile organic compounds (VOC) and solvents from the fermentation broths at much lower

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costs than distillation (Vane, 2005).

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Due to the higher potential for energy production and sustainability, the microbial conversion

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of biomass into valuable fuels and chemicals using syngas fermentation has been widely studied.

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Previous review reports (Abubackar et al., 2011; Daniell et al., 2012; Mohammadi et al., 2011;

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Munasinghe & Khanal, 2010a) have introduced the fundamental issues of syngas fermentation.

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However, poor GL-MT and kinetic limited conditions, which are still the primary barriers in the

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commercialization of syngas technology need to be summarized for the researchers working on

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the bioprocess design and scale up. This paper provides the rationales for syngas fermentation by

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considering the feedstock quality and process microbiology for the biological conversion of

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syngas. The fundamental aspects of GL-MT reactions are presented, and fermentation limited

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conditions are described. Various strategies to overcome each limitation are discussed, while an

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ideal case for optimum bioreactor operation is presented. Mass transfer performance and power

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consumption by various reactor configurations are summarized and compared. Finally, the

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overall challenges with each bioprocessing step are highlighted along with the potential

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strategies to resolve these problems. Future research directions are provided and a conceptual

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design for a hybrid membrane-based syngas biorefinery is proposed.

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2. Rationale for syngas fermentation

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2.1. Feed-stock quality

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Syngas is produced by the gasification of any type of organic matter from fossil and

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renewable resources, and is primarily comprised of CO, CO2 and H2 (Park et al., 2013). Trace

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amounts of NOx, SOx, tar, char, particulate matters (PM), C2 hydrocarbons (C2H2, C2H4 and

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C2H6), higher hydrocarbons (C6H6), NH3, HCN, chlorine compounds and water are also detected

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in the gasification process (Daniell et al., 2012). The components of syngas are found in various

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compositions depending upon type and quality of feed stocks, operating conditions in the gasifier

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and overall steps conducted for the gasification process. Gasification is fully matured and

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commercialized in the area of gas conversion and utilizing process. The final product

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composition from the gasification process is adjusted by changing the types of gasifier system

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(e.g., fixed bed, fluidized bed, etc.), using diverse feed stocks, and/or by maintaining the

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operating parameters (e.g., temperature, pressure, etc.) (Abubackar et al., 2011). The typical

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methods to restrict the composition of syngas to CO and H2 which are the primary components

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for biological fermentation process, involves the optimization of the gasification process

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conditions (Abubackar et al., 2011), bio-syngas production at temperatures of 1100°C (Kim &

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Chang, 2009), coal gasification at 1500-1800°C (Kim & Chang, 2009) and partial oxidation of

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biomass using oxygen/steam as the gasifying agents (Latif et al., 2014). N2 free syngas is

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produced using pure oxygen at the risk of high operating costs (Girard & Fallot, 2006).

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2.2. Process microbiology

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Microbes, mostly anaerobes, are used as biocatalysts to produce valuable metabolites, such

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as organic acids (acetic, butyric, formic acids) and alcohols (ethanol and butanol) from syngas

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fermentation (Table 1). These products are primarily end products during catabolic reaction

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indicating that they are produced during the ATP synthesis steps. Micro-organisms generally

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synthesize the ATP through, “substrate level phosphorylation (SLP)”, or “electron transport

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phosphorylation (ETP)”, in their energy metabolism (Latif et al., 2014). ATP synthesis through,

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an alternate mechanism, i.e., “chemical osmoisis” or “flavin-based electron bifurcation (FBEB)”,

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is verified in several acetogens (Latif et al., 2014). During FBEB, a key electron transfer agent of

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anaerobic microbes (i.e., reduced form of ferredoxin (Fd)), with high redox potential is

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generated, and is used for making ion motive force, termed as ATP synthesis. Difference in

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energy generation mechanism possibly leads to the diverse products formation in

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microorganisms. When syngas is used as sole carbon and energy source, acetyl-CoA is

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synthesized as a key intermediate, by the reductive acetyl-CoA (Wood-Ljungdahl) pathway,

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which has been hypothesized to anciently present in autotrophs before photosynthesis. CO and

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H2 are the energy sources for these microbes since 1 billion years before photosynthesis, which is

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believed to have evolved 3.8 billion years ago (Ragsdale & Pierce, 2008). During this energy-

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conserving process, CO2 is reduced as the terminal electron acceptor to acetyl-CoA, which is

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further reduced to create single and multi-carbon compounds, to consume reducing power.

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Anaerobic microorganisms (bacteria and archea) also produce energy carrier H2 via

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biological WGS reaction (Table 1). H2 production through the biological WGS reaction utilizes

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CO, which is the major component of the gasified syngas. Biological WGS for H2 production

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became a rising issue due to the versatile use of H2 as a primary energy carrier material and

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higher specific energy than coal and petroleum (Demirel, 2012), and the relatively fast CO

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conversion rates under thermophilic conditions for H2 production. Thermophiles do not only

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have higher growth and metabolic rates than mesophiles but they are also less prone to

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contamination, which is a major issue in pure culture fermentation processes. The

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hyperthermophilic archaeon, Thermococcus onnurineus NA1, isolated from a deep sea thermal

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vent, has unique hydrogenases and formate dehydrogenases, which overcome thermodynamic

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limitations of CO conversion to H2 when a proton is used as an electron acceptor under high

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temperature (Kim et al., 2010b).

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3. Bioprocess scale-up: from bench to industry

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The last decade of research has focused on CO fermentation, and fundamentals of syngas

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conversion pathways are well understood and communicated. A paradigm shift of research

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interests from fundamentals to practical applications in syngas/CO based biorefinery has

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occurred. Current and future research on syngas fermentation should be focussed to overcome

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the primary hurdles in the bioprocess scale-up for commercialization. Poor gas-liquid mass

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transfer, low achievable cell concentrations, toxic substrate inhibition and economical product

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separation are the real bottle necks that need to be addressed for the successful

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commercialization of the syngas fermentation technology.

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3.1. Gas-liquid mass transfer reaction

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Biorefinery of syngas components (CO, CO2 and H2) is a heterogeneous system

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consisting of gaseous substrate, liquid fermentation media and solid cells. The mass transfer of

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the substrate from the gas bubble to the reaction site in a cell is a complex process involving a

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series of resistances at a micro scale (Doran, 1995). However, syngas mass transfer through these

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resistances is considered to be controlled by only the liquid film across the gas-liquid interface

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(RL) (Klasson et al., 1992; Munasinghe & Khanal, 2010a; Vega et al., 1989b). All of the other

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resistances are minor; therefore, they can be neglected. The overall volumetric rate of mass

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transfer (R) of the sparingly soluble gaseous substrates through liquid film is the function of

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driving force (∆P or ∆C) times kLa as depicted by equation 4 (Table 2). Table 2 shows the

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leading mass balance equations that are used to represent the delivery of gaseous substrates to

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the liquid fermentation media. The relationship between the substrates delivery rate and

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substrates consumption under biotic conditions is also provided. Reducing the RL to create high

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dissolved substrate conditions inside the bioreactor is highly desired because the microbes can

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uptake the gaseous substrates only in liquid form. RL is effectively reduced by either increasing

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the driving force or kLa as depicted in equation 5.

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3.2 Bioreactor operation and fermentation limited conditions

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Design of a suitable bioreactor system is a core step for the successful operation of the

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syngas conversion bioprocess. The suitability of a proposed reactor system for GL-MT

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applications or absorption is founded on its capability to match with the reaction kinetics, and the

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mass transfer capabilities of these bioreactors are determined by a key parameter, kLa.

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Bioreactor operations for gaseous systems should be considered to be in either of two regimes in accordance with equations 13 and 14 (Klasson et al., 1992). The first case is the mass

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transfer-limited regime; which occurs when the cell mass in the system is greater than the solute,

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but the mass-transfer rate cannot keep pace supporting the microbial activity; therefore, the

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solute concentration in the liquid phase decreases. In the other case, kinetically limited regime,

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the high mass transfer reactor system delivers sufficient gaseous substrate to the microorganisms,

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but the lower cell concentrations in the fermentation media do not provide equal consumption

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rates. As a result, the concentration of the gaseous substrate in the fermentation media reaches

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the saturation value, making the conditions of no driving force creation for the mass transfer. In

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addition, higher concentrations of solute show a substrate inhibitory effect on the microbes,

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causing a decrease in the specific substrate consumption rate,
 . CO is a substrate that shows a

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substrate inhibitory effect above a critical concentration because CO inhibits metalloenzymes by

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forming stable complexes, resulting in reduced enzyme activity (Ragsdale, 2004). Notably, most

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enzymes in the acetyl-CoA pathway possess redox activities due to their metallic centers (Evans,

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2005). The optimal bioreactor designs permits high GL-MT and high cell concentrations under

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no substrate limiting and inhibitory conditions.

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3.2.1. Mass transfer limitation as a major problem and solutions to overcome this limitation

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Higher productivities of a bioprocess are strongly dependent on the high mass transfer

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rate and high cell concentration (Liu et al., 2014). Low mass transfer rates of gaseous substrates

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in the fermentation media lead to reduced cell concentrations (equation 7). The gas-liquid mass

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transfer was the most limiting factor in the syngas fermentation reaction, leading to reduced

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productivity (Abubackar et al., 2011; Munasinghe & Khanal, 2010a). CO conversion

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performance of Carboxydothermus hydrogenoforans in a gas lift reactor (GLR) was limited by

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the low cell concentrations (Haddad et al., 2014). Thus, in order to achieve an optimum

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bioprocess design, the reaction system should not experience the mass transfer limitations.

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Various strategies have been adopted to increase the GL-MT rates. The economic and practical

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viability of these approaches should be evaluated for their applications in large scale syngas

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fermentation systems.

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Bioreactors that achieve high mass transfer rates and high cell concentrations are desirable

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for syngas fermentation (Klasson et al., 1992). The highest kLa for GL-MT reactions have been

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achieved by using different reactor configurations (Orgill et al., 2013). Mechanical agitation used

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in STR is the most common method that is applied to enhance gas-liquid, liquid-liquid, and

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liquid-solid mixing in the industrial reactors. Maximum kLa achieved by mechanical agitation for

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GL-MT systems are summarized in Table 3a.

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Although continuous stirred tank reactors (CSTRs) are widely used, they do have

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intrinsic limitations that make them unsuitable for many microbial applications using a gaseous

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substrate. A high level of power-to-volume ratio in the agitator and rotational speeds of the

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impellers to enhance the mass transfer between the substrate and the microbes is not

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economically feasible for large reactors being considered for commercial syngas fermentations,

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primarily because of excessive power costs (Bredwell et al., 1999). The degree of agitation

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required to provide sufficient GL-MT for microbial growth may also damage sensitive

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microorganisms (Munasinghe & Khanal, 2012). CSTRs also have operating limitations such as

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biomass wash out at short hydraulic retention times and achievement of low biomass

272

concentration in gas system (Wang & Wan, 2009). Stirred tanks often exhibit poor mixing in

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tanks with large diameters and in multiphase reactions where the non-uniformity in mixing and

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mass transfer leads to significant variance in reaction rate and selectivity.

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Efforts have been made to achieve high GL-MT without agitation using TBR and BCR/GLR as shown in Table 3b. These types of reactors do not require mechanical agitation;

277

therefore, the power consumption of these reactors is lower than the STR. Bubble columns or

278

gas-lift systems are commonly used in industrial processes, both as reactors or absorbers,

279

whenever a large liquid retention time and/or a large liquid holdup are needed.

280

The achievable range of kLa using microbubble spargers (MBS) (200-1800 h-1) is much

281

higher than that of STR (10-500 h-1) (Bredwell et al., 1999). More than 40% of the research

282

conducted in the last decade (2005-2014) focused on the application of microporous hollow fiber

283

membranes (HFM) as the microbubble-generating system to achieve high GL-MT in syngas

284

fermentation (Table 4). The results suggested that the HFMBR may be the best alternatives of

285

agitation-based reactors for use in high GL-MT applications.

286

3.2.1.1. HFMBR

287

Recent research studies on the HFMBR have suggested that the MBRs have the potential

288

to replace the energy intensive, agitation-based systems used for high GL-MT applications. One

289

of the significant and recognized benefits of using the HFMBR as gas diffuser is its low or even

290

no direct energy consumption as the substrate gas can be pressurized directly to the reactor

291

through the storage cylinders. The working principle of HFMBR for GL-MT applications is

292

similar to membrane contactors where two phases are brought into contact to promote mass

293

transfer (Buonomenna, 2013). In syngas fermentation, the gas-liquid contact can occur either by

294

using inside-out or outside-in reactor configurations. In inside-out configuration, the gaseous

295

substrate is pressurized through the lumen of the HFM and diffuses directly into the fermentation

296

media across the wall of the membrane (Lee et al., 2012b; Orgill et al., 2013; Yasin et al., 2014).

297

Regarding outside-in configuration, the gaseous feed is supplied from the shell side of the HFM

298

and diffuses into the fermentation media circulated throughout the HFM lumen (Munasinghe &

299

Khanal, 2012). The choice of a suitable HFM has its own importance regardless of the type of

300

reactor configuration used. Hydrophobic membrane materials such as polypropylene (PP),

301

polyethylene (PE), and polyvinylidene fluoride (PVDF) are preferred compared with hydrophilic

302

materials due to their inherent advantages such as no pore wetting, less prone to biofouling and

303

ease in gas permeation (Yasin et al., 2014). The use of both porous and non-porous membrane

304

structure is possible to achieve high GL-MT. Non-porous structures are more stable and efficient

305

when high pressurized conditions are desired. The fabrication of hollow fiber membrane module

306

(HFMM) is another aspect that needs to be considered along with the membrane material and the

307

HFMBR configuration. Membrane modules are fabricated either in dead-end or open-end

308

configurations. An open-end configuration, however, leads to the higher loss of substrate gas.

309

Performance of HFMBR in terms of kLa outcompete all other reactor configurations

310

(Table 4). Shen et al. (2014b) achieved a kLa of 1096.2 h-1 using PP hollow fiber membranes.

311

This study has utilized Clostridium carboxidivorans P7 for biological conversion of syngas and

312

produced an ethanol concentration of 23.93 g L-1 with an ethanol to acetic acid ratio of 4.79.

313

Yasin et al. (2014) recently developed a simple high mass transfer HFMBR configuration

314

capable of supporting microbial conversion. This study introduced a novel idea to achieve high

315

driving force inside the reactor at low substrate pressures through the HFM lumen by increasing

316

the headspace pressure inside of the reactor.

317

3.2.2. Intrinsic kinetics as a substrate limitation and inhibition mechanism

318

Inhibition in the fermentation media is typically caused by the toxic nature of the

319

substrates and/or by the products and other contaminants. As a result, the specific microbial

320

growth rate (µ),
 and product formation rate are slowed or inhibited. Substrate inhibition

321

occurs during the initial oxidation of the electron donor substrates (Rittmann & MeCarty, 2005)

322

and reduces or stops the oxidation of the electron donor either by competitive or non-competitive

323

inhibition, resulting in slow substrate consumption rates (Rittmann & MeCarty, 2005). Several

324

microorganisms, such as “acetogens” that grow chemoautrophically on CO, are inhibited by high

325

dissolved concentrations of CO. CO has a high affinity to metal ions and potentially binds with

326

the cofactors of the CO-oxidizing enzymes and inhibits metalloenzymes by forming irreversible

327

complexes, which reduces their maximal activity (Ragsdale, 2004). The Andrews model (Table

328

2: equation 15 and 16) which is the modified form of the Monod equation, is used to determine

329

the intrinsic kinetic parameters, Ks (saturation constant) and Ki (inhibition constant), which are

330

further utilized to identify optimum concentrations of dissolved substrates in the fermentation

331

media using   (Andrews, 1968). The optimal dissolved CO concentrations (CL) for the

332

fermentation of P. productus was reported to be equivalent to 0.8 atm of partial pressure of CO

333

(PCO) using the non-completive substrate inhibition model proposed by Andrews (Vega et al.,

334

1988). Later Chang et al. (2001) and Younesi et al. (2005) found the maximal specific CO

335



uptakes (
 ) by Eubacterium limosum KIST612 and Clostridium ljungdahlii, respectively and

336

confirmed the non-competitive substrate inhibition in CO fermentation. The optimum CL to

337



obtain the maximum specific growth rate (µ max) and
 by E. limosum KIST612 was 1atm

338



(Chang et al., 2001), whereas an ideal activity of C. ljungdahlii in terms of
 was retarded

339

after 2 mM CL (Younesi et al., 2005). High pressure fermentation is detrimental when CL value is

340

greater than the kinetic requirements of the microbes; therefore, identifying the optimum CL to

341

design and control large scale fermenters is necessary. Substrate inhibition can be avoided by using a high mass transfer system capable of

342 343

supporting high cell density cultures, and high cell concentrations can be achieved in a

344

continuous cell recycled system (CCRS). The highest reported concentrations of ethanol (48 g L-

345

1

346

concentrations of 4 g L-1 (Phillips et al., 1993) were much higher than the levels produced in a

347

continuous syngas fermentation system without cell recycle (6 g L-1 ethanol at 2.3 g L-1 of cell

) that were achieved during syngas fermentation by C. ljungdahlii using a CCRS under cell

348

concentration) (Mohammadi et al., 2012). This finding showed the potential of using CCRS to

349

achieve high cell and product concentrations. Clostridium ragsdalei fermentation in a two-stage

350

CCRS yielded 15 g ethanol per g of cells (Kundiyana et al., 2011), which is comparable to 12 g

351

ethanol per g of cells obtained by Phillips et al. (1993) with C. ljungdahlii. A CCRS using mixed

352

culture and dominant with Alaklibacterium bacchi strain CP15 produced 6 g L-1 of ethanol (Liu

353

et al., 2014). However, the maximum attainable cell mass and product yields using CCRS

354

depends on the type of microorganism, the syngas composition and operating conditions of the

355

bioreactor. For example, the CO fermentation by E. limosum KIST612 only produced 0.3 g L-1

356

ethanol in a CCRS with a 4 g L-1 cell mass (Chang et al., 2001). Excessive loss of the nutrients

357

during a CCRS also needs careful economic evaluation.

358 359

3.2.3. Membrane biofilm reactor (MBfR): A potential alternative to resolve kinetic and mass transfer limitations

360

In conventional HFMBfR, the substrate gas is pressurized through the lumen of the fibers

361

and diffused directly in to the fermentation media through the walls of the membrane. Microbial

362

communities grow and attach to the membrane as aggregates and flocks resulting in the

363

formation of biofilm on the outer surface of hollow fibers, which are immersed in the

364

fermentation media. Here, HFM plays a bi-functional role as a gas diffuser and support for

365

microbial biofilm formation. Once the biofilm is fully developed, the gas diffused through the

366

membrane is taken up by the microbes within the biofilm and yields fermentation products such

367

as acitic acid and ethanol (Munasinghe & Khanal, 2010a).

368

MBfRs are a quite mature technology for water and wastewater treatment applications,

369

and one commercial scale facility, ARONiteTM has already been developed (Martin &

370

Nerenberg, 2012); yet, their applications in syngas fermentation for the transfer of CO, H2 and

371

CO2 are rarely found. One study has been reported regarding the production of medium chain

372

fatty acids in MBfR using H2 and CO2 as the substrates (Zhang et al., 2013). This study has used

373

the concept of microbial biofilm development using a mixed culture from a methane production

374

reactor operated under mesophilic conditions. The results showed 100% H2 utilization within

375

0.012 g biofilm, after 24 hours of continuous operation. Shen et al. (2014b) used a continuous

376

HFMBfR for syngas fermentation of C. carboxidivorans P7 and produced the highest

377

concentrations and productivity of ethanol (3.44 g per L-day) using the same strain. A recently

378

reported study also utilized an HFMBfR for the production of H2 from CO using a pure culture

379

of Clostridium hydrogenoformans (Zhao et al., 2013). The continuous uptake of CO within the

380

biofilm resulted in very low concentrations of dissolved CO in the liquid, improving the

381

permeation of the colonized membrane by three orders of magnitude as compared with the

382

abiotic membrane. The volumetric activity of C. hydrogenoformans in a GLR with suspended

383

growth was found to be 18-fold lower than the activity of a biofilm-based system. The outcome

384

of the above mentioned studies regarding HFMBfR have successfully demonstrated the

385

superiority of a biofilm (attached growth) reactor over suspended growth reactors.

386

4. Ideal bioreactor operation

387

In general, bioreactors adhering perfect mixing conditions with even distribution of nutrients

388

and dissolved substrates are expected to produce homogeneous productivities throughout the

389

reactor. However, real fermentation conditions may experience substrate limitation and

390

inhibitory conditions simultaneously in different parts of the reactor (Figure 1, left), depending

391

upon the availability of the dissolved substrate concentrations inside the reactor (Figure 1, right).

392

Microbial syngas conversion should follow the Monod kinetics under non-substrate limitation

393

and non-inhibitory conditions. However, CO is a well-known self-inhibitory substrate that

394

inhibits the specific microbial growth and substrate consumption rates in a non-competitive

395

manner. Therefore, a CO-fed bioreactor for microbial fermentation would be operated in either

396

of three of the following possible scenarios as shown in Figure 1 (right): substrate limited region

397

A ( <   ), ideal substrate region ( =   ) and substrate inhibitory region B

398

( >   ). For an ideal bioreactor operation to achieve maximum productivity, the

399

dissolved CO inside the reactor maintained at   should be continuously consumed.

400

However, these conditions may lead to region B due to dissolved substrate accumulation when

401

the microbial community does not provide equal consumption rates. The bioreactor operation in

402

region B is more likely to produce reduced productivities along with additional substrate loss.

403

The presence of higher amounts of substrate gas in the product also causes high downstream

404

separation costs. The most economically optimum operating regime for bioreactors for syngas

405

fermentation corresponds to the conditions where  ≤   . However, microorganisms under

406

this regime may experience the substrate limitation near the top of the bioreactor.

407

5. Overall challenges, strategies to resolve and future research directions

408

5.1. Microbial catalysts

409

The successful development and commercialization of a syngas biorefinery require the

410

development of robust microorganisms capable of producing higher yields of valuable chemicals,

411

materials and fuels. One of the major challenges associated with the existing syngas conversion

412

microorganisms is multiple product formation and lower yields (Table 1). Lower product

413

concentrations in the fermentation broth increase the products recovery costs, while the

414

production of undesired products causes additional costs for desired product separation and the

415

management of undesired products. To address these issues “Synthetic biology techniques” and

416

“metabolic engineering tools” have become very hot issues and several approaches have been

417

conducted to create recombinant strains to show strain development. These approaches are aimed

418

at the metabolic pathway shift for the production of desired chemicals and to improve the

419

robustness of the microbial strains. Genetic modification methods are applied to the microbial

420

strains in order to: accelerate the substrate consumption and the metabolic flux (I), increase the

421

productivity (II), produce more reduced form of product e.g., ethanol and butanol (III) and,

422

produce desired, non-native products by introducing new bio-circuit from foreign DNA.

423

Successful genetic transformation of Moorella thermoacetica has recently been reported for the

424

production of non-native products using genetic engineering tools (Kita et al., 2013). Among

425

acetogens, C. ljungdahlii, an ethanol producer, was genetically modified for butanol production

426

using a gene from C. acetobutylicum (Nagarajan et al., 2013). Genetic modification of

427

Clostridium autoethanogenum resulted in the production of 26 mM butanol (Liew et al., 2013), a

428

non-native product. Alcohol tolerance of this strain was enhanced by introducing alcohol

429

tolerance genes in the wild-type strain (Liew et al., 2013). Insertion of the genes related to

430

acetone synthesis, in Clostridium aceticum, resulted in the production of 140 µM acetone, a non-

431

native product (Liew et al., 2013). Higher concentrations of acetate were produced by

432

Acetobacterium woodii, after insertion of genes related to CO2 fixation and acetate synthesis, in

433

the wild-type strain (Straub et al., 2014). Genetically engineered, T. onnurineus NA1(MC01),

434

showed 30-fold higher transcription of the mRNA encoding CODH, hydrogenase, and Na+/H+

435

antiporter (Kim et al., 2013). Specific activity of H2 production on CO, by MC01 was 1.8 fold

436

compared with wild type strain, while H2 production potential was boosted 3.8 fold compared

437

with wild-type strain. These developments have paved the way for desired product formation

438

from industrial waste gases by the genetic modification of existing acetogens. However, this

439

approach is quite time and labor intensive and requires comprehensive metabolic information.

440

An extensive research work is required to develop accurate genome-scale metabolic models,

441

preferably models incorporating regulatory mechanisms (Lee et al., 2012a). General metabolic

442

engineering tools that can be employed to engineer any strain of interest are also desired.

443

Long term adaptation is recently reported as an effective method to; understand the

444

mechanisms of molecular evolution, and to obtain desired features in the microbial strains, such

445

as; growth on the desired substrates, and increase in nutrients and environmental stress resistance

446

(Dragosits & Mattanovich, 2013). Clostridium thermoaceticum growth on heterotrophic

447

substrate (i.e., saccharide) was shifted to autotrophic growth (100% CO or H2/CO2), to show

448

establishment of the acetyl-CoA pathway by adaptation (Kerby & Zeikus, 1983). Pre-adaptation

449

of C. ragsdalei on producer gas showed improvement in tolerance to syngas impurities. The pre-

450

adapted strain showed 22% higher final cell mass concentration, and 1.9 and 2.8 fold enhanced

451

ethanol productivity, using producer gas as substrate (Ramachandriya et al., 2013).

452

5.2 Bioprocess design

453

Poor mass transfer of CO and H2 is a key bottle neck in the scale-up of syngas

454

fermentation technology. The problem has been realized and research has been conducted to

455

investigate various reactor configurations to support high mass transfer rates. The last decade has

456

witnessed a paradigm shift in research focus from conventional agitation-based reactors (Table

457

3a) to membrane based reactors (Table 4). Membrane-based reactors have outcompeted

458

conventional reactors in terms of GL-MT and have significant economic advantages compared

459

with agitation-based reactors. Power consumption, which is based on methods to increase GL-

460

MT and the scaleup issues of various reactor configurations, are summarized in Table 5.

461

In addition to many advantages, the application of HFMBR in real fermentation facilities

462

is hindered by several factors. Yasin et al. (2014) highlighted the key research perspectives for

463

using HFMBR for high GL-MT applications. Challenges with HFMBR starts with the selection

464

of proper membrane materials. So for the research conducted in syngas fermentation using HFM

465

has utilized the hollow fibers that were primarily made for waste water treatment applications.

466

Thermal stability and durability of these hollow fibers for GL-MT need to be assessed. The

467

commercial fabrication of HFM for syngas fermentation has business potential, and the

468

fabrication of HFM specifically designed to achieve high GL-MT of syngas components would

469

offer better mass transfer and durability. Submerged HFMBR configurations with dead ends are

470

economically more viable. However, the effect of bio-fouling and short gas retention times

471

require careful consideration before commercial applications. HFMs have shown the highest GL-

472

MT efficiency, and future research should be focused on their practicality under real

473

fermentation conditions. Long term trials with actual fermentation broths are desired to assess

474

the membrane and module stability and fouling factor, and comparing the performance of

475

suspended growth reactors with biofilm reactors will be interesting.

476

CO is an inhibitory substrate; therefore, it is desired to find intrinsic kinetics of the

477

microorganism that would be used in the syngas conversion process. An ideal bioreactor should

478

match the intrinsic requirement of the microorganism without substrate limitation or inhibition.

479

The desired mass transfer rate to achieve optimum substrate levels in bioreactors is achieved by

480

controlling kLa or the driving force for the mass transfer (C*− CL) (equation 4). Achieving a

481

high driving force of the gas in a bioreactor is difficult; therefore, studies have historically

482

focussed on obtaining a high kLa. Overall, the mass transfer rates to the bioreactors should be

483

sufficient to support maximum cell growth (equation 7); while an ideal bioreactor should be

484

operated in accordance with equation 9.

485 486

The use of high cell density cultures to achieve high productivities should also be reevaluated in economics terms. The current approaches for using high cell density cultures result

487

in the excessive loss of the nutrients and also require extra pumping costs for the addition of

488

fresh medium and the removal of permeate.

489

5.3. Product separation

490

A successful biorefinery require a hybrid continuous fermentation system (HCFS)

491

involving the in situ separation and recovery of fermentation products. The selection of an

492

optimal product separation and purification process will play a critical role in the scaleup and

493

commercialization of fermentation facilities. The fermentation of gasified syngas requires a

494

suitable process for the removal of fatty acids (e.g., formic and acetic acids) and VOC (ethanol,

495

butanol and isopropanol). The continuous removal of products from the fermentation broths may

496

increase the productivity of fermentation especially when the syngas-converting microbes exhibit

497

inhibition either by fatty acids or solvents. Distillation has been the dominant technology for the

498

recovery of VOCs from fermentation broths; however, product separation and recovery at a

499

smaller scale are not economical, particularly when the fermentation broths have lower

500

concentrations of the desired products. Table 6 provides the technological options that can be

501

integrated into a syngas based-biorefinery for the recovery and separation of desired products.

502

The choice of the desired separation technology depends on several factors. Economic feasibility

503

and practical viability should be assessed at various biofuel production scales before choosing

504

any technology. Membrane pervaporation appears to be a promising future technology due to the

505

simplicity of operation, ease in the scale-up and scale-down and relatively low cost. A

506

conceptual design of hybrid multiple membrane modules installed in an HFMBR system is

507

presented in Figure 2. The issue of low residence time associated with stand-alone HFMBR

508

configurations may be addressed by using multiple membrane modules. In addition, the possible

509

un-homogenized conditions at high cell concentrations may be avoided by media circulation,

510

which is primarily conduced for product separation. However, the success of membrane

511

pervaporation highly depends upon the development of ideal membranes for multiple products

512

separation. The manufacturing and use of highly selective membranes for specific products may

513

allow a single vessel product recovery, which reduce the cost for infra-structure and downstream

514

processing.

515

6. Conclusions

516

The vision for the sustainable production of biofuels and chemicals can be realized by the

517

successful scale-up of syngas fermentation. Overall, the challenges with this technology are

518

highlighted together with the potential solutions by considering all process components. Special

519

focus is given to the bioprocess design and HFMBRs are proposed to resolve the issues of GL-

520

MT and the kinetic limitations. A hybrid bioprocess that combines the upstream, midstream and

521

downstream technologies by the application of synthetic biology, metabolic engineering,

522

bioprocess engineering and membrane technology is expected to pave the way for economical

523

and efficient utilization of organic matter from fossils and renewable resources.

524

Acknowledgements

525

We would like to thank Engr. Abdul Waheed Bhutto (Dawood University, Karachi) and Dr.

526

Aqeel Ahmed Bazmi (COMSATS University, Lahore) for their valuable comments. This work

527

was supported by the New & Renewable Energy Core Technology Program of the Korea

528

Institute of Energy Technology Evaluation and Planning (KETEP) granted financial resource

529

from the Ministry of Trade, Industry & Energy (No. 20133030000090), and the Development of

530

Biohydrogen Production Technology Using the Hyperthermophilic Archaea Program of the

531

Ministry of Oceans and Fisheries in the Republic of Korea.

532

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51. Phillips, J.R., Klasson, K.T., Clausen, E.C., Gaddy, J.L., 1993. Biological production of ethanol from coal synthesis gas. Appl. Biochem. Biotechnol. 39-40, 559-571. 52. Ragsdale, S.W., 2004. Life with Carbon Monoxide. Crit. Rev. Biochem. Mol. Biol. 39, 165195. 53. Ragsdale, S.W., Pierce, E., 2008. Acetogenesis and the Wood–Ljungdahl pathway of CO2 fixation. BBA-proteins Proteom. 1784, 1873-1898. 54. Ramachandriya, K., Wilkins, M., Patil, K., 2013. Influence of switchgrass generated producer gas pre-adaptation on growth and product distribution of Clostridium ragsdalei. Biotechnol. Bioprocess Eng. 18, 1201-1209. 55. Ramaswamy, S., Huang, H., Ramarao, B. 2013. Separation and purification technologies in biorefineries. Wiley. 56. Riggs, S.S., Heindel, T.J., 2006. Measuring carbon monoxide gas-liquid mass transfer in a stirred tank reactor for syngas fermentation. Biotechnol. Prog. 22, 903-906. 57. Rittmann, B.E., MeCarty, P.L. 2005. Environmental technology: principles and applications. McGraw-Hill. 58. Robinson, C.W., Wilke, C.R., 1973. Oxygen absorption in stirred tanks: A correlation for ionic strength effects. Biotechnol. Bioeng. 15, 755-782. 59. Shen, Y., Brown, R., Wen, Z., 2014a. Enhancing mass transfer and ethanol production in syngas fermentation of Clostridium carboxidivorans P7 through a monolithic biofilm reactor. Applied Energy. 136, 68-76. 60. Shen, Y., Brown, R., Wen, Z., 2014b. Syngas fermentation of Clostridium carboxidivoran P7 in a hollow fiber membrane biofilm reactor: evaluating the mass transfer coefficient and ethanol production performance. Biochem. Eng. J. 85, 21-29. 61. Sim, J.H., Kamaruddin, A.H., Long, W.S., Najafpour, G., 2007. Clostridium aceticum-a potential organism in catalyzing carbon monoxide to acetic acid: application of response surface methodology. Enzyme Microb. Technol. 40, 1234-1243. 62. Straub, M., Demler, M., Weuster-Botz, D., Dürre, P., 2014. Selective enhancement of autotrophic acetate production with genetically modified Acetobacterium woodii. J. Biotechnol. 178, 67-72. 63. Ungerman, A.J., Heindel, T.J., 2007. Carbon monoxide mass transfer for syngas fermentation in a stirred tank reactor with dual impeller configurations. Biotechnol. Prog. 23, 613-620. 64. Vane, L.M., 2005. A review of pervaporation for product recovery from biomass fermentation processes. J. Chem. Technol. Biotechnol. 80, 603-629. 65. Vane, L.M., 2008. Separation technologies for the recovery and dehydration of alcohols from fermentation broths. Biofuels, Bioprod. Biorefin. 2, 553-588. 66. Vega, J., Holmberg, V., Clausen, E., Gaddy, J., 1988. Fermentation parameters of Peptostreptococcus productus on gaseous substrates (CO, H2/CO2). Arch. Microbiol. 151, 65-70. 67. Vega, J.L., Antorrena, G.M., Clausen, E.C., Gaddy, J.L., 1989a. Study of gaseous substrate fermentations: carbon monoxide conversion to acetate. 2. continuous culture. Biotechnol. Bioeng. 34, 785-793. 68. Vega, J.L., Clausen, E.C., Gaddy, J.L., 1989b. Study of gaseous substrate fermentations: carbon monoxide conversion to acetate. 1. batch culture. Biotechnol. Bioeng. 34, 774784. 69. Wang, J., Wan, W., 2009. Factors influencing fermentative hydrogen production: A review. Int. J. Hydrogen Energ. 34, 799-811.

710 711 712 713 714 715 716 717 718 719 720 721 722 723 724 725 726 727 728 729 730 731

70. Yagi, H., Yoshida, F., 1975. Enhancement factor for oxygen absorption into fermentation broth. Biotechnol. Bioeng. 17, 1083-1098. 71. Yasin, M., Park, S., Jeong, Y., Lee, E.Y., Lee, J., Chang, I.S., 2014. Effect of internal pressure and gas/liquid interface area on the CO mass transfer coefficient using hollow fibre membranes as a high mass transfer gas diffusing system for microbial syngas fermentation. Bioresour. Technol. 169, 637-643. 72. Younesi, H., Najafpour, G., Ku Ismail, K.S., Mohamed, A.R., Kamaruddin, A.H., 2008. Biohydrogen production in a continuous stirred tank bioreactor from synthesis gas by anaerobic photosynthetic bacterium: Rhodopirillum rubrum. Bioresour. Technol. 99, 2612-2619. 73. Younesi, H., Najafpour, G., Mohamed, A.R., 2005. Ethanol and acetate production from synthesis gas via fermentation processes using anaerobic bacterium, Clostridium ljungdahlii. Biochem. Eng. J. 27, 110-119. 74. Zhang, F., Ding, J., Zhang, Y., Chen, M., Ding, Z.-W., van Loosdrecht, M.C.M., Zeng, R.J., 2013. Fatty acids production from hydrogen and carbon dioxide by mixed culture in the membrane biofilm reactor. Water Res. 47, 6122-6129. 75. Zhao, Y., Haddad, M., Cimpoia, R., Liu, Z., Guiot, S.R., 2013. Performance of a Carboxydothermus hydrogenoformans-immobilizing membrane reactor for syngas upgrading into hydrogen. Int. J. Hydrogen Energ. 38, 2167-2175.

732

Figure Captions

733 734

Figure 1: Dissolved CO profile with reactor height (left) and three cases of bioreactor operation (right)

735

Figure 2: Multiple membrane modules installed HFMBR system integrated with pervaporation

736 737 738 739 740 741 742 743 744 745 746 747 748 749 750 751 752 753 754 755

756

757

Figure 1

758

759

Figure 2

760

Table 1: Selected syngas utilizing microorganisms Growth condition¬

Bacteria

Domain

Microorganisms

Type Isolated from

Mud

Yeast Temp. pH extract (℃ ) (g L-1)

Acetobacterium woodii

Wild

A. woodii pJIR750 THF

recombi nant

Butyribacterium methylotrophicum

Adapted Sewage sludge on CO digestor 7.3

Clostridium aceticum

Wild

Mud

7

30

2

7

30

4

Growth Parameters Substrate

Culture mode

Product

-1

OD/g L / no. of

-1

µ (h )

cells CO2-H2 = n/a –PH2 1700 Batch-STR mbar CO2-H2 =167-400 Batch-STR (mbar)

(mM)

(mM h-1)

Reference

n/a

acetate: 745.76

5.2

Yes

(Demler & Weuster-Botz, 2011)

1.2-2

0.056

acetate: 864.4

20.3

Yes

(Straub et al., 2014)

0.5

CO:CO2 = 70:30

Batch-serum vial

0.35

0.042

acetic acid: 21.6, acetic acid: 0.15, ethanol: 1.7, ethanol: 0.01, butyric acid: 4.5 butyric acid: 0.03

No

(Heiskanen et al., 2007)

2

CO:H2:Ar=78:4:18

Batch

0.8

n/a

acetate: 31.4

No

(Sim et al., 2007)

1

CO:CO2:H2:N2 =20:20:10:50

Batch

0.15

Yes

(Cotter et al., 2009)

4.5-6 37

0.5

CO:CO2:H2:N2 =20:15:5:60

ContinuousHFMBfR

1.7 (batch n/a mode)

Yes

(Shen et al., 2014b)

4-6

37

1

Batch-serum vial

1.15

0.022

Yes

(Younesi et al., 2005)

4.5

36

-

CO:CO2:H2:Ar =55:10:20:15 CO:CO2:H2:Ar =55:10:20:15

CCRS

4

n/a

Yes

(Phillips et al., 1993)

Yes

(Perez et al., 2013)

No

(Kundiyana et al., 2010)

Yes

(Chang et al., 2001)

7.18.9

37 30

Wild

Rabbit faeces 6

Clostridium Carboxidivorans P7

Wild

Agricultural lagoon

Clostridium ljungdahlii

Wild

Clostridium ljungdahlii

Wild

Clostridium ljungdahlii ER1-2

Wild

Natural water 4-5.5 35 source

0.5

CO:CO2:H2=60:5:35

Batch

OD6003

0.094

Clostridium ragsdalei (P11)

Wild

Duck pond sediment

5-6

37

-

CO:CO2:H2:N2 =20:15:5:60

Semicontinuous

1.13

n/a

Eubacterium limosum KIST612

Wild

Anaerobic digester

6.8

37

1

CO 100%

Batch-vial* Batch-BCR

1.35

0.17*

Moorella thermoacetica

Productivity

1.1

Clostridium autoethanogenum

Chicken yard waste

Con.

Whole genome sequencing

37

Wild Horse manure 7

0.5

CO2:H2= 20:80

Batch

OD660 0.18

n/a

acetate:26.9

0.32

Yes

3

CO 100%

Batch

OD660 1 n/a

acetate:58.9

0.45

Yes

0.16

0.2

acetate: 16.9

2.8

Yes

(Vega et al., 1989a)

0.106

0.1

YH2:83.3̂ acetate: 0.51 ethanol: 0.087 methanol 0.25

n/a

Yes

(Haddad et al., 2014)

0.7*

H2 : 171.3

5.71

No

(Jung et al., 2002)

n/a

acetate: 8.0

acetate:0.021

NO

(Küsel et al., 2000)

55

Moorella thermoacetica

Adapted on CO

Peptostreptococcus productus

Wild

Anaerobic digester sludge,

7-7.4 37

2

CO:CH4:CO2=63.43:20. CSTR 61 :15.96

Carboxydothermus hydrogenoformans

Wild

hot spring

6.8-7 70

0.05

CO 100%

Citrobacter sp Y19

Wild

Clostridium drakei

wild

Anaerobic WW 7 sludge digester Acidic 5.8-

30

n/a

acetate: 0.93

acetate: 23.3, acetate:0.35, ethanol: 1.45 ethanol: 0.02 acetate:108.4, acetate: 1.9, ethanol:519.4 ethanol: 3.1 butanol: 6.07 acetate: 20.3, acetate: 0.21, ethanol: 13.0 ethanol: 0.14 acetate: 50.8, Acetate:0.09, ethanol:1042 ethanol:1.86 acetate: 123.27, Acetate:0.58, ethanol 135.72: ethanol:0.45 ethanol: 548.30 ethanol: 0.39 acetic acid: 80.27 acetic acid: 0.14 1-butanol: 6.34 1-butanol: 0.02 acetate: 90, acetate: 1.38 butyrate: 0.7 butyrate: 0.01

3

25-30 0.5

Batchcontinuous CO supply

Batch-Serum CO:air vial* 2* +organics=2.5:97.5 (v/v) Batch-STR H2 Batch-infusion 0.06

(Kerby & Zeikus, 1983)

Archaea

sediment

Thermococcus onnurineus Wild NA1 Thermococcus onnurineus recombi NA1 nant (MC01)

761 762

¬

6.9

Hydrothermal 6.1vent 6.2

CO

bottles

CO 100%

BatchContinuous CO supply BatchContinuous CO supply

10 80 10

0.08

acetate: 7.3

acetate: 0.019

0.43

0.31

H2 : 332.2

31.8

Yes

(Kim et al., 2013)

1.62

0.72

H2 : 1641.6

123.5

Yes

(Kim et al., 2013)

Growth conditions used in this table are specific to the referred studies. The optimum growth conditions can be found elsewhere (Abubacker et al., 2011; Liew at al., 2013; Mohammadi et al., 2011; Munasinghe & Khanal, 2010). ; (̂ % mol mol-1 CO)

763

Table 2: Leading mass balance equations Equation   = −  (1)

Background/application ● Fick's first law of diffusion is the basis of mass transfer in binary (gas-liquid) phases i.e., two film theory by Whitman.



 =  ×  × ∆



● Applicable when ∆ = !∆ (Henry’s law).

(2)

 =  ×  × ∆ (3) "=



∆ '* ()

=

#

=  ( ∗ −  ) (4)

● Overall mass transfer rate of gaseous substrate to liquid medium.

'

(5)

● RL = resistance to the mass transfer caused by the liquid film.

0 = −( )1 (6)

● Solution of equation (3) under fully homogenized conditions. ● Used to find kLa in bioreactors without considering the effect of gas and/or liquid flow rates. ● At steady state, when there is no accumulation of gas (CO) in the reactor, the overall rate of CO transfer from the bubbles should be equal to the rate of CO consumption by the cells. ● Under biotic conditions, substrate uptake rat is equivalent to specific substrate consumption times cell concentration. ● Obtained by the comparison of equations (4) and (8). ● Represent ideal case for optimum bioreactor operation in the absence of substrate limitation and inhibition. ● Expression relating the QCO with cell concentration, specific growth rate and the cell growth yield coefficient.

+)

ln .1 −

)

∗

" =
 2 (7)
 2 = 4 (8)   ( ∗ −  ) =
 2 (9) 4 = μ= 6

(10)

78⁄9: 6

 ( ∗ −  ) (11)

= =6⁄  ( ∗ −  ) (12)

#

4 = −

56

78⁄9:

() A

D ' C9:

B)

μ= ) I9:

5

#

E



−

M

M

5FGH

) EFGH I9:

M

=

) I9:

M

EFGHJN  

+

+

) I9:

A

+

● Replacement of μ with

>? '

>@ ?

● Rate expression for CO uptake under mass transfer controlled conditions. ● Substrate transfer rate per unit working volume of the reactor incorporating the inhibitory effect of gaseous substrate. ● Modified Monod Model (MMM) equation proposed by Andrew’s for the determination of intrinsic kinetic parameters. ● Differentiation of equations 15 and 16 leads to equations 15a and 16a which can be used as second order non-linear regression models to find Ks and Ki.

JK

5FGH

(16)

EFGH

() .B)

(14)

(15)

) I9:

+

5FGH JN

) LI) *J JK LI9: N 9:

=

(13)

) ) 6EFGH I9:

M

) I9:

 )

) LI) *J JKLI9: N 9:

) 5FGH I9:

(15a)

) I9:

=

 

) LI ) *J JKLI9: N 9:

=


=

(

● Solution of equation (9) and (10).

+

D C9:

#

JK

EFGH

(16a)

(17)

● Relation used to find the partial pressure of CO in liquid phase in equilibrium with the gas phase.

764 765 766 767 768

Notations : a: interfacial area gas phase (m2 m-3), CA: concentration of a component A (mol m-3), C*: saturated gas concentration at equilibrium (mol m-3),

769 770

qCO: specific CO consumption rate (mol gcell-1 h-1), QCO: CO consumption rate per unit volume of broth (mol h-1 L-1), t: time, μ: specific growth rate (h-1), VL:

CL: the dissolved gas concentration (mol m-3), ∆C: concentration gradient (mol m-3), D AB: diffusivity of component A through B, which is a measure of its diffusive mobility (m2 s-1), H: Henry law constant (L atm mol -1), JA: molar flux of a component A relative to the average molal velocity of all constituents (mol m-2 s-1), kG and kL: diffusion coefficient of per unit area through the gas and liquid films respectively (m s-1), NA: mass transfer rate (mol m-3 s-1), NCO: moles of CO,

 :

partial pressure of CO in the liquid phase,

  :

partial pressure of CO in the gas phase, PM: particulate matter, ∆P: pressure gradient,

working volume of reactor (m3), =6⁄: growth yield coefficient (gcell produced (mol CO consumed)-1), z: distance (m) in z-direction.

771

Table 3a : Maximum k La in mechanical mixed (MM) reactors Reactor type

Hydrodynamic condition Substrate

Reference

(Klasson et al., 1993) (Bredwell et al., 1999)

CSTR ST-MBS

CO CO Syngas

Agitation speed (rpm) 700 200 300

STR

CO Syngas Syngas O2 Syngas

600 700 500 900 300

n/a n/a R. rubrum n/a SRB mixed culture

CO Syngas

200 300

B.methylotrophicum C. ljungdahlii

35.5 90.6 104 for CO & 190 for H2 154.8 292.68 72.8 114 31 for CO & 75 for H2 14.2 35 for CO

Syngas CO CO

450 150 300

R. rubrum n/a n/a

101 for CO 33.5-53.3 34.9-55.8

SpargerMM

Microorganism

kLa (h-1)

R. rubrum B.methylotrophicum SRB mixed culture

(Riggs & Heindel, 2006) (Kapic et al., 2006) (Younesi et al., 2008) (Orgill et al., 2013) (Bredwell et al., 1999)

(Munasinghe & Khanal, 2010b)

772 773

Table 3b: Maximum kLa in non-agitated membrane less reactors Reactor type

Hydrodynamic condition

kLa (h-1)

Reference

774 775

Substrate

Microorganism

776 BCR

CO CO CO CO O2 Syngas

n/a n/a n/a n/a n/a SRB MC

Syngas

R. rubrum

72 94.3 400 450 421 121 for CO & 335 for H2 55.5

Syngas CO Syngas CO

C .ljungdahlii R. rubrum R. rubrum n/a

137 38 2.1 2.5-40

CO

n/a

31.7-78.8

CO

n/a

29.5-50.4

785

GLR GLR

CO CO

n/a n/a

16.6-45.0 49.0-91.1

(combined with a single-point gas entry) 786 (combined with 20 µ m bulb diffuser)

GLR GLR GLR

CO H2 CO

n/a n/a n/a

129.6 97.2 *1.5-2

(Munasinghe & Khanal, 2014)

MlBfR TBR

(Chang et al., 2001) (Park et. al, 2013) (Shen et al., 2014a) (Orgill et al., 2013) (Bredwell et al., 1999)

777 778 779 780 781

PBR Column diffuser 20 µm bulb diffuser Sparger

(Cowger et al., 1992) (Bredwell et al., 1999) (Munasinghe & Khanal, 2010b)

782 783 784

787

790

(Haddad et al., 2014)

(MlBfR: monolithic biofilm reactor, PBR: packed bed reactor, *at 70°C, SRB: sulphate reducing bacteria, MC: mixed culture)

788 789

791

Table 4: Maximum k La in HFMBR Membrane characteristics

Configuration

As/VL (m-1)

Pressure through lumen (kPa)/Water recirculation rate (mL min-1)

Mixing reservoir (rpm)

kLa (h-1)

Reference

Membrane Material

Water interaction

Pore size (µm)

ID/OD (µm)

PVDF

Hydro Phobic (on CO)

0.1

700/1200

Stand alone1 (inside out)

62.5

37.23/Not used

Not used

155.2

(Yasin et al., 2014)

PP

Hydro Phobic (on CO)

0.04

220/300

External2 (inside out)

175

103.4/1000

200

1096.2

(Shen et al., 2014b)

PVDF

Hydro Phobic (on CO)

0.2

800/1400

Internal3 (inside out)

2250

203/1500

Not used

1.36^

(Zhao et al., 2013)

PS

Hydrophilic (on O2)

n/a

500/660

External (inside out)

4366

1-2SLPM*/80

Not used

55

(Orgill et al., 2013)

PES

Hydrophilic (on O2)

n/a

1100/1300

External (inside out)

2271

1-2SLPM*/80

Not used

23

(Orgill et al., 2013)

PDMS

Hydrophobic (on O2)

Nonporous

200/300

External (inside out)

10000

1-2SLPM*/400

Not used

1062.0

(Orgill et al., 2013)

PP

Hydro Phobic (on CO)

0.2

376/426

External (inside out)

56

114.5/670

90

385.01

(Lee et al., 2012b)

CHF

Hydro Phobic (on CO)

n/a

200/240

External (outside in)

200

206.8/1500

Not used

946.0

(Munasinghe & Khanal, 2012)

CHF

Hydro Phobic (on CO)

n/a

n/a

Internal (inside out)

200

241/500

Not used

1.08

(Munasinghe & Khanal, 2010b)

792

CHF: Composite hollow fiber, n/a: Not available , PE: Polyethylene, PP: Polypropylene, PS: Polystyrene, PES: Polyethersulfone, PDMS: Polydimethylsiloxane, PVDF: Polyvinylidenefluoride, SLPM:

793

standard liter sper minute, * It is given that gas inlet pressure was maintained between 0.7 and 4.8kPa to attain the desired flow rate, 1Submerged type HFMBR that act as a sole reactor without gas and

794

media circulation, 2CO diffuser is installed in the main reactor, 3CO diffuser is installed externally to the main reactor, ^at 70°C

795

796

797

Table 5: Economic and scaleup issues for different reactor configurations Reactor type

Ways to achieve high kLa

STR

● Increase in agitation. ● Increase in volumetric gas flow rate. ● Variation in vessel geometry.

BCR

● Substrate flow rate through pump.

TBR

● Gas flow rate.

HFMBR with media circulation

● Membrane surface area. ● Membrane material. ● Fluid (substrate and medium) circulation (flow rate) through pump. ● Substrate pressure.

Stand-alone HFMBR without media circulation

● Membrane surface area. ● Membrane material. ● Substrate pressure.

Power consumption equation ● Q = RS TURV WX Where, P= Un-gassed power consumption NP = empirical dimensionless power number n=number of impellers ρ= liquid density N=Agitation speed d= diameter of impeller (Davis, 2010). ● Pumping is required for fluid (gas or liquid) circulation in BCR, TBR and HFMBR with media circulation YZ[ Q= \]]] Where, P= required pumping power in kw ^= Specific weight of the fluid = ρ * g (N m-3) Q=fluid recycle rate (m3 s-1) h= Pressure head loss through the system (m) (Kim et al., 2010a).

● No pumping and agitation are required.

Economics

Scaleup issues

● A linear increase in kLa is possible by increasing the impeller speed. ● The power required for a STR is proportional to the cube of impeller speed, thus a linear increase in mass transfer will cause cubic increase in power consumption.

● Increase in gas flow rates cause the reduction in gas utilization due to small retention times in liquid which leads to reduced conversion efficiencies. ● Not suitable for shear sensitive microorganisms.

● About a linear increase in k La is possible by increasing the substrate flow rates. ● Power consumption increases linearly with gas flow rates. ● kLa mainly depends upon the gas flow rate (Orgill et al., 2013). ● A linear increase in power consumption for a linear increase in gas or liquid flow rates.

● Tall reactors shell will be required to achieve 100% substrate conversion in the fermentation medium. ● Maintaining the homogeneity may be difficult at higher cell concentrations. ● Complex fluid dynamics. ● Higher gas flow rates yields to lower retention times, substrate loss and lower productivities.

● kLa can be increased by increasing the membrane surface area, inlet gas flow rates and liquid media flow rates. (Lee et al., 2012b; Orgill et al., 2013; Shen et al., 2014b). ● Power consumption increases linearly with increase in flow rates of substrate gas and fermentation medium. ● No additional power is required for fluid circulation. Substrate gas can be supplied directly through pressurized gas cylinders (Yasin et al., 2014).

● Unavailability of commercial hollow fiber membranes for mass transfer applications. ● Additional power consumption through liquid media circulation. ● Possible cell washout due to medium circulation at higher velocities.

● Lower residence times of substrate gas in the fermentation medium which eventually leads to reduced productivities. ● Possibility of high amount of waste gas productions. ● Maintaining homogeneity may be difficult at higher cell concentrations.

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Table 6: Technological options for product recovery and separation in syngas fermentation. Extracted from (Ramaswamy et al., 2013; Vane, 2008) Technology

Working principle/ Application area ● Cause the separation of different components from feed stream based on differences in their boiling points. ● Can be used for solvent recovery.

Advantages

Limitations

● Offer ≥90% alcohol recovery. ● Easy to scaleup. ● Enough energy efficiency at moderate feed concentrations (>4 wt% ethanol).

Gas stripping

● The VOC produced during fermentation can be stripped by gases and then condensed to leave the concentrated solvent solutions. ● Can be used for solvent recovery.

Liquid-liquid extraction (LLE) or Solvent extraction

● Solvents are separated based on differences in their solubility in extractants and the aqueous phase for separation. ● Can be used for the recovery of solvents and fatty acids.

● Can selectively remove the desired product. ● Do not harm any microorganism. ● Possible to use an undesired fermentation byproduct (CO2) as stripping gas. ● In-situ product recovery. ● Well understood mechanism. ● Well-established chemical operation.

● Scale down is difficult. ● Require costly product dehydration. ● High operating temperature is harmful for many microbes. ● Inadequate in the presence of azeotropes (e.g. water and ethanol). ● Cause excessive foaming of the fermentation broth. ● Additional step for solvent recovery. ● Useful for only large volumes of fermentation broths.

Distillation

Adsorption

Membrane Pervaporation

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● Desired chemical compounds are transferred from liquid streams to a adsorptive materials (adsorbents). ● Can be used for recovery of solvents and fatty acids.

● Usually the solid adsorbents are stable and do not dissolve in liquid components of fermentation broth.

● Two or more miscible components are separated from a liquid stream by vaporization, based on their different affinities for membrane material and difference in diffusion rates through the membrane. ● Can be used for recovery of solvents and fatty acids.

● Required energy is much lower than distillation. ● Can be scaled down due to simplicity of operation ● Successfully applied ‘‘in situ’’ during acetone, butanol and ethanol (ABE) fermentation (de Vrije et al., 2013).

● Emulsion formation. ● Low distribution coefficients for conventional extractants. ● Highly selective solvent is required. ● Requires either high temperature or pressure, which is detrimental for active microorganisms. ● Complexity of scale-up in biorefining processes. ● Requires desorption of adsorbent. ● Possibility of undesired adsorption phenomenon in real fermentation conditions. ● Biofilm formation on the surface of adsorbing particles. ● Less selective separation. ● Possible increase in temperature due to evaporation of liquid stream. ● Require costly product dehydration. ● An ideal membrane can limit the transport of desired specie under certain scenarios (Vane, 2005).

Technical and economic viability ● Fully developed technology. ● Uneconomical at smaller scale. ● Net energy obtained from ethanol < energy required for distillation at lower ethanol concentrations in feed streams i.e., < 1 wt% (Vane, 2005). ● High separation costs for the removal of VOC from gas stream. ● High cost of the compressor when used for industrial applications. ● Well-developed infrastructure. ● Higher costs for extractants. ● Costly regeneration of extractants. ● No available infra-structure for biological conversion reactions.

Qualitative comparison

● Require low energy. ● Energy intensive regeneration of adsorbate. ● High adsorbent cost.

● Low capacity ● Low selectivity ● High fouling ● Easy operation

● Relatively immature technology. ● Low energy input. ● Energy required to evaporate and condense the undesired specie could be substantially high compared to the energy required to evaporate and condense the desired specie.

● Moderate capacity ● Moderate selectivity ● Low fouling ● Easy operation

● High capacity ● High selectivity ● No fouling ● Complex operation

● Moderate capacity ● Low selectivity ● Low fouling ● Easy operation

● High capacity ● High selectivity ● Moderate fouling ● Difficulty in operation

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Highlights

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Rationales for syngas bioconversion are provided. Strategies to resolve kinetic and mass transfer limitations are presented. An ideal case of bioreactor operation for syngas bioconversion is proposed. Mass transfer and economic performance by various reactor configurations is compared. A conceptual design for a hybrid membrane-based syngas biorefinery is proposed.