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Systems Microbiology and Biomanufacturing (2024) 4:365–385 
https://doi.org/10.1007/s43393-023-00205-z
REVIEW
Strategies and engineering aspects on the scale‑up of bioreactors 
for different bioprocesses
Ariane Fátima Murawski de Mello1  · Luciana Porto de Souza Vandenberghe1  · 
Leonardo Wedderhoff Herrmann1  · Luiz Alberto Júnior Letti1  · Walter José Martinez Burgos1  · 
Thamarys Scapini1  · Maria Clara Manzoki1  · Priscilla Zwiercheczewski de Oliveira1  · Carlos Ricardo Soccol1 
Received: 28 July 2023 / Revised: 11 September 2023 / Accepted: 13 September 2023 / Published online: 10 October 2023 
© Jiangnan University 2023
Abstract
Bioreactors are central equipment used in the majority of bioprocesses. Different models of bioreactors have been devel-
oped for different processes, which can be applied either for submerged or for solid-state fermentation. Scale-up involves 
the development of bioprocess in bench, pilot, and industrial scales. Optimal conditions are first screened and determined 
in the bench scale and so that the process can be transferred to a larger scale. This transferring requires the proper reproduc-
tion of conditions and performance, being a major challenge since important aspects, such as aeration and agitation, are 
critical for cells development. In this case, scale-up strategies are employed to maintain bioprocesses’ performance. These 
strategies are based on geometric similarity aspects of bioreactors, agitation, and aeration conditions, which must follow the 
requirements of each bioprocess and the used microorganisms. Operational conditions significantly impact cell growth and, 
consequently, the biosynthesis of different biomolecules, which must then be reproduced at higher scales. For this purpose, 
one or more operating factors can be maintained constant during scale-up with the possibility to predict, for example, the 
power consumption of large-scale bioreactors or aeration conditions in an aerobic culture. This review presents the most 
employed bioreactors’ scale-up strategies. In addition, the scale-up of other bioreactors models, such as pneumatic and solid-
state fermentation bioreactor and even photobioreactors, will also be described with some examples.
Keywords Bioreactors · Bioprocess · Bioproducts · Scale-up · Biomolecules
Introduction
Bioprocesses and fermentation have been developed since 
antiquity, with products such as beer, wine, bread, vinegar, 
and others, being obtained with the purpose of food con-
servation [1]. With technology development and knowledge 
about how these bioprocesses occur, highly pure biomol-
ecules could be produced in larger scales. Nowadays, it is 
possible to manufacture a variety of bioproducts—such as 
organic acids, biofuels, bioplastics, biopesticides, pharma-
ceutics, aromas, and others—by fermentation. However, this 
kind of process is usually more complex than the chemical 
routes, mainly because cells with specific physicochemical 
requirements are used [2]. Besides, in order for a bioprod-
uct to reach the market, some steps need to be followed for 
proper process implementation and scale-up.
Usually, bioprocess development will occur in three 
scales: laboratory or bench, pilot and, finally, industrial 
(see e-Supplementary material) [3]. In bench scale, small 
volumes are applied, either in flasks (50–500 mL) or in 
bioreactors (1–15 L) for conditions screening and process 
optimization. Therefore, different nutrients sources and 
physicochemical conditions (such as temperature, pH, agi-
tation, and aeration rates) for maximum biomass and product 
yields and minimum cost can be easily tested [4]. Besides, 
optimization and simulations via experimental design can 
be conducted, and models can be developed and validated 
on this scale [5]. With all the optimal conditions determined 
and the process tested in a bench-scale bioreactor, it can be 
transferred to pilot scale (bioreactors of 50–500 L).
This process will implicate in maintaining some param-
eters (mainly involving aeration and agitation) constant 
 * Luciana Porto de Souza Vandenberghe 
 lvandenberghe@ufpr.br
1 Department of Bioprocess Engineering and Biotechnology, 
Federal University of Paraná, Centro Politécnico, Curitiba, 
Paraná 81531-980, Brazil
http://orcid.org/0000-0003-4103-8462
http://orcid.org/0000-0003-0267-1185
http://orcid.org/0000-0002-2301-3942
http://orcid.org/0000-0003-3733-6133
http://orcid.org/0000-0003-4377-6530
http://orcid.org/0000-0003-1184-3049
http://orcid.org/0000-0002-0892-8132
http://orcid.org/0000-0002-2688-8090
http://orcid.org/0000-0001-7630-6864
http://crossmark.crossref.org/dialog/?doi=10.1007/s43393-023-00205-z&domain=pdf
366 Systems Microbiology and Biomanufacturing (2024) 4:365–385
1 3
among scales, along with the geometrical similarity of the 
bioreactors. The pilot scale will function as a demonstration 
step for determining if the developed bioprocess is viable 
and establishing important parameters that could not be opti-
mized in the laboratory (e.g., agitation influence in shear 
forces) [6]. With the economic and technical viability of 
the project being demonstrated in pilot scale, the bioprocess 
can finally reach the industrial and commercial steps. Scal-
ing up the process from pilot to industrial will also involve 
similarity criteria to assure the process success. Bioprocess 
scale-up will, therefore, always involve the design of bio-
reactors among all scales. This type of equipment can be 
considered the heart of the bioprocesses as they will hold 
the cells needed for the bioproduct manufacturing. There-
fore, the accurate choice of the bioreactor is imperative in 
all scales, depending on the mode of operation, moisture 
content, financial resources available, and cells applied [7].
In this sense, bioprocess and bioreactor scale-up are 
straightly interconnected. During scale-up, it is important 
to provide the precise conditions for microbial development 
and growth, while guaranteeing the economic and technical 
viability of the project. In the present review, the aspects 
influencing scale-up, along with the main strategies applied 
for different bioreactors, will be discussed and examples of 
implemented industrial processes and bioreactors will be 
given.
Bioprocess aspects influencing scale‑up
Operation modes
Bioprocesses can be operated in different manners along 
the fermentation time, depending on the microbial demands 
in terms of substrate consumption, product formation, and 
possible inhibitions. Essentially, there are three modes of 
operation (Table 1): batch, fed-batch, and continuous, and 
they differ accordingly to feeding of fresh media and/or 
withdrawal of fermented broth during the process [8]. The 
simplest mode of operation is batch, wherein there is no 
addition or withdrawal of media across the fermentation 
time. Therefore, due to metabolic dynamics, the broth is 
constantly changing, generating an unsteady state [9, 10]. 
Batch operation mode can be highly applied in a labora-
tory scale for initial tests of production, and screening the 
optimal conditions for producing the desired biomolecule. 
Similar to batch fermentation, there is no withdrawal of fer-
mented broth during the fed-batch process. However, there 
is addition of new media during the fermentation time [11]. 
The feeding solution can consist only of the carbon source, 
or of a nutrient solution, or it can be the complete media, 
depending on the nutritional needs of the cells and it can Ta
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atsimilar. According, to 
the authors, 60–75% of the CO2 introduced in the PBR was 
mitigated by the microalgae, while 25–40% of the CO2 was 
exhausted from the PBR to the atmosphere, meaning a total 
of 535 kg of CO2 consumed to produce 296 kg of biomass 
in the 100 m3 PBR during a 60-day operation. Besides the 
good CO2 mitigation efficiency, photosynthetic efficiencies 
of up to 3.5% of total solar irradiance were attained, as well 
as promising biomass and lipid productivities—with the pos-
sibility of many biotechnological applications [144].
Chlorella vulgaris is a microalgae species remarkable for 
its versatility, presenting high lipid contents and significant 
amounts of vitamins, minerals, proteins, antioxidants, and 
pigments. It can be used for producing high-value chemi-
cals, cosmetics, and pharmaceuticals, since it presents anti-
oxidants, anticancer, antimicrobial, antidiabetic, antihyper-
tensive, and antihyperlipidemic activities [145]. It can also 
be sold directly as food supplement in the form of powder, 
extracts, capsules, or tablets [145, 146].
Chlorella vulgaris tolerates high CO2 concentrations, 
showing good mitigation rates with reasonable growth 
[146]. In the study from Paladino and Leviani [146], Chlo-
rella vulgaris was cultivated in glycerol rich wastewater 
and CO2, in airlift photobioreactors whose scale-up was 
based on Buckingham π-theorem. In industrial scaling-up, 
besides considering the increase in work volumes, it is still 
essential to consider the changes in operational mode and 
in bioreactor type. In this practical case, initially the 
microalgae were cultivated in STR, and were scaled up to 
airlift PBRs, which allows proper photoperiods and good 
mixing without high energy demands. The π-theorem was 
used to define the main 12 dimensionless numbers, called 
π numbers (such as Re, Sh, ds
d
,
T
Topt
,
Io
Kl
, pH, etc.), at lab scale 
and to keep their values as desired at pilot scale.
Mass transport, global kinetics, and dimensionless 
numbers adopted to perform scale-up were obtained from 
the 0.5 L DSTRs to semi-continuous 2.5 L STRs by exper-
imental campaigns. To further scale up from semi-contin-
uous 2.5 L STRs to semi-continuous 10 L airlift reactors 
(ALRs), a combination of approaches was employed, cou-
pling fluid dynamics experimentation. Finally, scale-up 
verification at pilot-scale ALRs was performed by comput-
ing from the experimental campaign in outdoor conditions 
the remaining dimensionless numbers related to the kinet-
ics of algae growth and process yield. These computed 
numbers aligned with the expected values based on the 
previous results obtained from the 0.5 L DSTRs, demon-
strating the feasibility of scaling up microalgae cultivation 
in PBRs using the π-theorem [146].
Aligned with the biorefinery concept and circular 
approaches, microalgae can be utilized in wastewater treat-
ments. Liquid agro-industry wastes are generated in enor-
mous amounts, normally having significant concentrations 
of nutrients, and through microalgae cultivation it is pos-
sible to aggregate value at the same time that COD values 
are reduced and water reusing is enabled [147]. Dairy liq-
uid effluents are one example of agro-industrial wastewater 
that can be treated by using microalgae. In the article from 
Kumar et al. [148], high-volume V-shape Ponds (HVVP) 
were proposed to establish higher volume to surface ratios 
and lower land foot-print compared to the conventional 
microalgae open raceway ponds (ORPs). HVVP is a 
V-shaped channel-like structure, specific for phycoreme-
diation of industrial effluents, and notably cheaper com-
pared to vertical or horizontal tubular photobioreactors 
[148]. The pilot-scale V-shape ponds have a size of 2 × 2 m 
(occupying an area of 4 m2), and a maximum working 
volume of 3 m3, for a depth of 1 m. The inverted pyramid 
shape provides a maximum surface area to the microalgae 
for light absorption (S/V ratio of 1.33). Aeration is pro-
vided through interconnected PVC pipes located at the 
bottom of the pond and at its half the height, guaranteeing 
uniform circulation and exposure of the microalgae cells 
to the light. With the results in pilot scale for the micro-
algae Ascochloris sp. ADW007, an economical study was 
performed about projected scenario cases: for treatment 
capacity plants of 0.25, 0.5 and 1.0 million liters of dairy 
effluent generated per day, showing that HVVP is found to 
381Systems Microbiology and Biomanufacturing (2024) 4:365–385 
1 3
be one of the cost-effective and area-efficient microalgal 
cultivation systems for mass production [148].
Research needs and future prospects
The growing concern about human activity in the environ-
ment has led to the rise of commercial bioprocess and bio-
products as potential alternatives to the conventional ones 
with the development of biofuels, bioplastics, alternative 
food and feed, among others. Although these bioprocesses 
share some similarities with their chemical counterparts, 
there are some specificities that differentiate them, such 
as the need of proper agitation and aeration for proper cell 
development. As explored throughout this review, different 
bioreactors can serve as vessels that support microbial cells 
and bioproducts formation depending on the bioprocess 
conditions and requirements. Scaling-up and commercial-
izing the final product still remains a challenge. Therefore, 
there is an urgent need for technical and economical analysis 
of the developed processes in order to identify gaps prior 
to scaling-up. Besides, this review showed different strate-
gies for scaling-up distinct types of bioreactors that can be 
adapted to several bioprocesses. To guaranteeing the com-
mercial success of the developed product, researchers are 
encouraged to test their processes both in bench and pilot 
bioreactors, being able to screen conditions that directly 
affect cell development. With the constant development of 
new bioproducts, new models of highly technological bio-
reactors can also be proposed.
Conclusions
A bioprocess begins at bench scale, where the process’s 
conditions are defined and optimized. However, the defined 
conditions must be transferred to larger scales (pilot and 
industrial scales). The success of a process transfer depends 
on the correct choice of scale-up strategies, which are based 
on important parameters, such as agitation and/or aeration, 
which must be maintained at the new scale. Each process 
presents its peculiarities, having some specific exigences 
with a perfect combination of the binomial aeration- agi-
tation, promoting optimal microbial growth and efficient 
production of the desired bioproduct. It is also important to 
choose the correct bioreactor model and mode of operation 
and define the combination of one or more scale-up crite-
ria to achieve better process performances. It is clear that 
efforts have been made to modify and/or adapt the known 
design of submerged and non-submerged bioreactors. Even 
if basic designs of bioreactors remain the same, new studies 
for their modification and scale-up are continuously being 
carried out, trying to respond to the recent evolution of the 
biotechnology industry.
Supplementary Information The online version contains supplemen-
tary material available at https:// doi. org/ 10. 1007/ s43393- 023- 00205-z.
Acknowledgements The authors thank the Coordenação de Aper-
feiçoamento de Pessoal de Nível Superior (CAPES) and Conselho 
Nacional de Desenvolvimento Científico e Tecnológico do Brasil 
(CNPq) for the Project fundings and research scholarship
Author contributions AFMdM conceptualization, writing—original 
draft, writing—review. LPdSV conceptualization, writing—original 
draft, writing—review. LWH conceptualization, writing—original 
draft. LAJL conceptualization, writing—original draft. WJMB writ-
ing—original draft. TS writing—original draft. MCM writing—origi-
nal draft. PZdO writing—original draft. CRS projectadministration, 
funding acquisition.
Funding Projects funding and research scholarship are provided 
by Coordenação de Aperfeiçoamento de Pessoal de Nível Superior 
(CAPES) and Conselho Nacional de Desenvolvimento Científico e 
Tecnológico (CNPq).
Availability of data and materials Not applicable.
Declarations 
Conflict of interest The authors declare that they have no known com-
peting financial interests or personal relationships that could have ap-
peared to influence the work reported in this paper.
Ethical approval and consent to participate Not applicable.
Consent for publication All authors have read and agreed to publish 
the final version of the manuscript.
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	Strategies and engineering aspects on the scale-up of bioreactors for different bioprocesses
	Abstract
	Introduction
	Bioprocess aspects influencing scale-up
	Operation modes
	Water and solid content strategies
	Factors affecting bioprocess scale-up
	Agitation, aeration, and viscosity
	Similarity criteria
	Scale-up strategies for submerged processes
	Stirred tank reactor
	Constancy of volumetric mass transfer coefficient (kLa)
	Constancy of power input per volume of medium (PV)
	Constancy in Reynolds number (NRe) and mixing time (tm)
	Pneumatic
	Photobioreactors
	Scale-up strategies for solid-state processes
	Bioprocess scale-up examples
	Biofuels
	Food and ingredients
	Waste treatment
	Bioactive compounds
	Microalgae bioprocess
	Research needs and future prospects
	Conclusions
	Acknowledgements 
	Referencesio
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367Systems Microbiology and Biomanufacturing (2024) 4:365–385 
1 3
be fed into the bioreactor in pulses (intermittent) or in a 
continuous mode [11].
The main objective of the fed-batch is to prolong the log 
phase of the cell development, achieving therefore higher 
biomass and product yields [12]. The determination of the 
time to start feeding can be tricky and needs to be well inves-
tigated and researched. Usually, feeding starts when the car-
bon source has reached a certain percentage of consumption 
or has been completely depleted from the media, or when 
microbial development has reached its maximum, aiming 
therefore to prolong this stage. The feeding needs to be 
done in a well-established strategy to avoid media and cells 
dilution, which will hinder the final yield of the bioprocess 
[8, 13]. Another factor that is determinant in fed-batch is 
the control of process parameters. The feeding of exter-
nal nutrients will impact in several fermentation aspects: 
concentration of biomass, product and substrate, dissolved 
oxygen, growth rate, among others [14]. Open- and closed-
loop control systems are the most common types of control 
used in fermentation processes; however, new strategies are 
constantly being developed and researched to guarantee the 
success of the fed-batch [12, 14–16].
In the continuous mode of operation, there is both feed-
ing of fresh media and removal of fermented broth from the 
bioreactor, maintaining a steady state internally (chemostat) 
and, generally, a constant volume throughout the fermenta-
tion process [17]. Immobilized cells in different supports 
can be applied in these cases with more ease, avoiding wash-
outs—a phenomenon that happens when cells are unable to 
grow faster than they are removed [18, 19].
Water and solid content strategies
Water content is a determinant factor in bioprocesses, as it 
will impact not only on how the cells propagate, but also in 
the choice of the type of bioreactor (Table 2). Solid-state 
fermentation (SSF) is the type of fermentation where the 
water content is low (between 30 and 85%), therefore occur-
ring in a solid support [28]. It is a method that can be applied 
mainly for the cultivation mainly of fungi that will grow 
through the solid matrix [29]. However, other microorgan-
isms, such as Bacillus sp., can also propagate in solid sub-
strates [30]. Agroindustrial substrates can be directly applied 
in these systems as they serve as the solid support for the 
microbial development [31]. The main challenge involving 
SSF is precisely the scale-up of this type of process. Unlike 
the submerged fermentation, the relations between bioreac-
tors’ dimensions and shapes are not so easy to establish and 
some parameters involving water flow and velocity cannot 
be applied for scaling-up [32].
On the other hand, the semi-solid-state fermentation is 
the type of fermentation that contains a high content of 
water, but has solids in suspension during the fermentation. 
This fermentation overcomes the challenges of mass and 
heat transferring of the SSF. Most of the bioreactors that are 
applied to the submerged fermentation can also be applied 
to the semi-solid-state fermentation [31]. The submerged 
fermentation (SmF) is the most common type of fermenta-
tion applied in the industry. Microbial growth and devel-
opment and the biomolecules production occur in a liquid 
media with water activity above 0.95 [33]. The definition of 
which system and conditions that will be applied will rely 
on process parameters developed in bench scale and in what 
is economically feasible.
Factors affecting bioprocess scale‑up
Agitation, aeration, and viscosity
Most industrial bioprocesses by submerged fermentation 
are aerobic in aqueous medium enriched with macro and 
micronutrients. Generally, these broths are viscous and 
behave like non-Newtonian fluids. In these processes, 
Table 2 Different fermentation strategies and their characteristics
Parameter Solid-state fermentation Semi-solid-state fermentation Submerged fermentation References
Characteristics Fermentation occur in a solid 
matrix, water activity under 
0.95
Substrate solids are suspended in 
a liquid broth, minimum water 
activity of 0.95
Fermentation occur in a liquid 
broth, minimum water activity 
of 0.95
[31]
Advantages Low chance of contamination, 
high product concentration, 
easy downstream
Easy handling, controlling and 
scaling-up
Easy handling, controlling and 
scaling-up
[31, 34, 35]
Disadvantages Difficulty in scaling-up and con-
trolling, mass and heat transfer 
can be compromised
Difficulty on downstream, chance 
of contamination, cleaning 
issues
Difficulty on downstream, chance 
of contamination
[31, 35]
Most applied bioreactors Packed-bed, trays, rotary drum, 
fluidized-bed bioreactor
Stirred tank reactor, pneumatic 
(airlift and bubble column)
Stirred tank reactor, pneumatic 
(airlift and bubble column)
[36, 37]
Potential biomolecules Enzymes, organic acids, tradi-
tional food and medicine
Bioethanol, enzymes, biopesti-
cides
Ethanol, enzymes, bioplastics, 
antibiotics, among others
[38, 39]
368 Systems Microbiology and Biomanufacturing (2024) 4:365–385
1 3
oxygen is essential for the growth and maintenance of 
microorganisms, as well as for the production of the bio-
metabolite, since oxygen is used as an electron acceptor 
[34]. Therefore, the supply of oxygen to the system must 
be ensured and the transfer of oxygen in the broth must 
be known. Aeration is a challenge during an industrial 
aerobic process, as oxygen has low solubility in water and 
needs to overcome several diffusion barriers in order to 
reach the microbial cells [35]. Therefore, these factors 
need to be taken into account while scaling-up. When 
applying bioreactors, different agitation and aeration sys-
tems, along with viscosity controlling along fermentation 
time, can be applied to provide the needed oxygen to the 
cells, efficient mass and heat transferring, and bioreac-
tor homogeneity. Agitation will be strictly dependent on 
the type of bioreactor applied. While stirred tank reactors 
(STRs) provide agitation through mechanical stirring with 
impellers [36], pneumatic bioreactors can homogenize the 
media with bubbles (in the case of bubble column) or air 
flux (in the case of airlift) [37]. Cell type needs to be taken 
into account, as some microorganisms can be more sensi-
ble to high agitation rates than others [7].
Aeration is usually provided by a system of air com-
pressor (which will pump the air for the bioreactor), an air 
cooler (which will chill the air temperature if necessary) 
and a sparger (which will help to distribute the air in the 
bioreactor along with the agitation system) [40]. The rate 
of aeration should be determined by the oxygenuptake rate 
(OUR) that sets how much oxygen is being consumed by 
the microorganisms along time [41]. The oxygen transfer 
rate (OTR), on the other hand, is related to the oxygen con-
centration that passes through the media along time, and is 
directly related to the mass transfer coefficient (kLa) [42, 43]. 
Therefore, OTR and kLa can be used as scaling-up criteria 
in order to guarantee a great aeration for bioreactors sys-
tems. Agitation and aeration can imply foam formation dur-
ing fermentation time [44]. Besides, some bioprocess will 
form more foam than others due bioproducts characteristics 
(e.g., biosurfactant production) [45]. Therefore, as foam can 
impact in air diffusion, there is an imperative need for proper 
headspace dimensioning when designing the bioreactors, 
and the installation of a foam formation control system with 
anti-foam substances being added when necessary [46].
Viscosity is another parameter that can influence aeration, 
as an increase in media viscosity can influence in the OTR, 
along with the bubble coalescence and distribution [47]. Vis-
cosity can be caused by solids presence (e.g., in semi-solid-
state fermentation), microbial biomass development, and 
product formation (especially when the product is viscous, 
as xanthan gum). In these cases, when the fermented broth 
becomes more viscous along time, agitation and aeration 
become critical factors that need optimization in laboratory 
and pilot scale prior to industrial implementation. Studying 
several impellers configuration, and different aeration rates 
is needed for the success of the bioprocess [48].
Similarity criteria
Apart from all the relations established for the scaling up of 
bioprocess, at the end of the day, scaling-up ultimately relies 
on the similarity. The developed bioprocess and/or bioreac-
tor will have some conditions, parameters, relations and/or 
ratios that need to be equal along all the scales for the pro-
cess to be reproducible and viable, therefore accomplishing 
the same objective which is the production of a determined 
biomolecule in high yield and productivity [3, 49]. Thus, 
not only is reinforced the need of a well-established and 
optimized process in bench scale, it is aso needed that these 
conditions persist while scaling up. Obviously, it is not pos-
sible to maintain all the characteristics between scales, but 
some of them are indispensable. These similarities can be: 
chemical and biochemical, mechanical, thermal, and geo-
metrical [49].
The main geometrical ratio that has to be maintained 
along scales is the ratio between the height (H) and the 
diameter of the vessel (D), which vary with the type of the 
bioreactor (e.g., STRSs have a H:D ratio of 1:1 while bubble 
columns can reach up to 10:1) [3]. Generally, a higher H:D 
ratio will imply a higher heterogeneity in the bioreactor as 
agitation can be compromised in different points (top and 
bottom). Therefore, this ratio needs to be taken into account 
when designing a tall bioreactor (in the case of a STR, it is 
needed that the impellers can cover all the points in the sys-
tem, varying the number and/or distance of impellers) [50]. 
Other relations that can be determined are among the height 
liquid (HL) and the reactor (HR), that needs to be in between 
0.7 and 0.8 for proper headspace and foam formation, and 
relations between impeller (DI), baffles (Db) and bioreac-
tor (Dt) diameters in STRs [11, 50, 51]. These geometric 
relations can be really useful in scaling up, and it is impor-
tant to establish these criteria. However, other biochemical 
parameters (e.g., physicochemical characteristics, microbial 
growth rates) also need to be taken into account while scal-
ing up applying multifactorial analysis for proper process 
reproduction [3, 52].
Scale‑up strategies for submerged processes
Stirred tank reactor
The stirred tank reactors (STRs) are the most used types of 
reactors or bioreactors in the bioprocess industry and indus-
try in general. These bioreactors are mainly composed of a 
tank, which is provided with an agitation system with one 
or more impellers mounted on a shaft. In addition, these 
369Systems Microbiology and Biomanufacturing (2024) 4:365–385 
1 3
systems can also be provided with other components, such 
as sprinklers, baffles, sensors, coils and suction and supply 
pipes (Fig. 1) [36, 53]. It is necessary to know the dynamics 
and the quantitative relationships between the parameters 
of the bioreactor, as well as their influence on the type of 
metabolite that will be produced. The processes for scaling 
up the production of bioproducts are complex because not 
all the parameters established on the laboratory or bench 
scale can be maintained on the larger scale, which is mainly 
due to the fact that the parameters or scaling criteria are 
interrelated, so that the change of one parameter can affect 
another one. Therefore, in the scaling-up processes, it is nec-
essary to evaluate the most significant parameters that must 
be maintained in the largest scale. The main criteria or scale-
up factors most used in STR are: constancy of volumetric 
oxygen transfer coefficient (kLa) and potency per unit vol-
ume of medium (P/V) [54]. Other less used scaling criteria 
or parameters are: constancy in Reynolds number, constancy 
in mixing time (tm), constancy of velocity at the impeller, 
and constancy of impeller pumping capacity.
Constancy of volumetric mass transfer coefficient (kLa)
Generally, kLa is used as a scale-up criteria in bioprocesses 
that demand large amounts of oxygen, such as the production 
of antibiotics [54, 55], exopolysaccharides production [56], 
and recombinant proteins [57]. To use the kLa constancy 
criteria, it is established the fact that kLa is proportional to 
the power transmitted to the fluid under aeration, the volume 
of the medium and the superficial velocity (Eq. 1). These 
correlations are valid for Newtonian fluids.
where kLa: volumetric transfer coefficient of O2 (h−1), Pg: 
power transmitted to the fluid over aeration (W), V: medium 
volume (m3), Vs: surface speed (m. s−1), Vs is given by the 
Eq. 2.
where Qs: air volume flow (m3 s−1), S: cross-sectional area 
of the tank (m2) (π DT
2/4), DT: tank diameter.
Assuming kLa constancy between scale 1 and 2 (Eq. 3):
Therefore, Eqs.  1 and 3 are combined, considering 
that the power in a non-aerated system can be related to 
the power in a gasified system (Pg) through the correla-
tion of Michel and Miller [58]. In addition, it should be 
considered that the volume is proportional to the cube of 
impeller diameter (Di), power is proportional to impel-
ler diameter multiplied by agitation velocity, and sur-
face velocity (Vs) is proportional to air volume flow (Qs) 
divided by the square of Di. After these considerations, it 
is obtained that to keep the kLa constant, it is necessary 
(1)kLa ∝
(
Pg
V
)A
.
(
Vs
)B
,
(2)Vs =
Q
s
=
(
4Q
�D2
T
)
,
(3)(kLa)1 = (kLa)2.
Fig. 1 STR model with geometric similarity relations. Adapted from [11, 54]
370 Systems Microbiology and Biomanufacturing (2024) 4:365–385
1 3
to relate the diameters of the impellers of the bioreactors, 
as well as the flow rate (Q) of the two scales (Eq. 4). In 
fact, Bandaiphet and Prasertsan [56] reported that kLa is 
significantly affected by system geometry and other oper-
ating parameters such as system impeller speed [59, 60]. 
In addition, coefficients A and B must also be considered, 
which depend on the volume of the bioreactor (Table 3).
where N: rotating frequency (rps or s−1), Di: impeller diam-
eter (m), Qs: air volume flow (m3 s−1);
Shin and collaborators [61] scaled the production of 
itaconic acid from 5 to 50L, using as a scale-up criterion 
the constancy of the kLa parameter (0.02  s−1). The biomass 
and acid production in the smaller scale were 12 g/L and 
51.2 g/L, respectively, and on the larger scale they were 
12.2 g/L and 52.7 g/L. Furthermore, the specific growthrate (µ) for 5L and 50L were 0.029  h−1 and 0.031  h−1, 
respectively. The obtained results showed that the kLa 
constancy is an excellent scaling strategy for the produc-
tion of itaconic acid since there is no significant difference 
between the yields obtained in the two scales.
Constancy of power input per volume of medium (P/V)
The constancy of power per unit volume (P/V) is another 
widely used criteria in bioprocess scaling-up. In general, 
this parameter is used for scaling the production of alco-
hols, organic acids and mammalian cell cultures [3, 62], 
that is, processes that are not aerated or where the trans-
fer of oxygen does not turn out to be so significant. It is 
known that in bioreactors in cylindrical tanks with baffles 
and stirred by impellers in laminar and transition regimes, 
the power number (NP) is an inverse function of the Reyn-
olds number or modulus (NRe) (Eq. 5) [54].
where NP and NRe are dimensionless numbers, and are 
expressed by Eqs. 6 and 7, respectively.
(4)N2 = N1
(
Di2
Di1
)
2B−2.85A
3.15A
.
(
Q2
Q1
)
0.25A−B
3.15A
,
(5)NP = f
(
1
NRe
)
,
In scale-up processes, the physical properties of the fluid 
remain constant; therefore, the density and viscosity of the 
culture medium are also constant. Furthermore, in the turbu-
lent regime NP is also a constant. In the scaling process, the 
first strategy is to maintain the geometric similarity, so the 
volume is proportional to the impeller diameter (Di). After 
these considerations, the expression for scale expansion is 
obtained, keeping the criterion (P/V) constant (Eq. 9). While 
keeping the parameter (P/V) constant, impellers’ diameters 
and speed need to be well evaluated as some microorganisms 
can be sensible to high agitation rates.
where P: power (W); V: medium volume (m3); N: impeller 
speed (s−1); Di: impeller diameter (m).
Constancy in Reynolds number (NRe) and mixing time (tm)
The NRe and tm parameters are rarely used scaling criteria 
because they are directly linked to the degree or speed of 
agitation (N). Therefore, selecting some of these parameters 
as the main criteria indicates a possible change in the param-
eters kLa or (P/V) which may affect microbial performance. 
Considering that NRe1 = NRe2 and that in the scaling pro-
cesses the physical properties of the fluid remain constant, 
therefore the NRe (Eq. 7) is reduced to a simple proportion-
ality (NRe α NDi
2). The expression for scaling keeping NRe 
constant is shown in Eq. 10, which shows a relationship 
between impeller diameter (Di) and speed (N) [54].
The mixing time (tm) can be defined as the time required 
for fluid homogenization [54]. To obtain rapid mixing, the 
bioreactor must have a robust agitation system. However, the 
tm is affected by the properties of the fluid, so when the fluids 
used are viscous or non-Newtonian the tm increases along 
with the required power [3]. The mixing time factor (Φ) is 
related to the NRe. According to Norwood and Metzner [63] 
and Schmidell et al. [54], the mixing time factor is an inverse 
function of NRe. Considering that NRe > 105, Φ reaches a 
(6)NP =
P
N3D5
i
�
,
(7)NRe =
ND2
i
�
�
.
(8)
(
P
V
)
1
=
(
P
V
)
2
,
(9)N2 = N1.
(
Di1
Di2
)2∕3
,
(10)N2 = N1
(
Di1
Di2
)2
.
Table 3 Values for coefficients 
A and B for different volumes 
of bioreactors
Source [3].
Volume (m3) A B
0.005 0.95 0.67
0.5 0.6–0.7 0.67
50 0.4–0.5 0.5
0.002–2.6 0.4 0.5
371Systems Microbiology and Biomanufacturing (2024) 4:365–385 
1 3
constant value of approximately 4. Under these conditions 
of NRe and considering that HL and DT are proportional to 
Di, tm can be expressed as shown in Eq. 11. Considering 
tm1 = tm2 the final expression for scaling is shown in Eq. 12.
where tm: mixing time (s), N: impeller speed (s−1); Di: impel-
ler diameter (m); HL: fluid column height (m); DT: tank 
diameter (m).
Pneumatic
Pneumatic bioreactors are characterized by homogenization 
and agitation processes through gas bubbling in the reac-
tion vessel, being the most common bubble column (BCR) 
and airlift bioreactors (Fig. 2) [3]. The scale-up processes 
for pneumatic reactors follow the premises of STRs, and 
it is essential to observe criteria such as heat, mass, and 
flow transport phenomena (mainly related to aspersion and 
gas flow), mixture characteristics, geometrical similarities, 
and reaction kinetics [3, 64]. In addition, in pneumatic reac-
tor scalability designs, there are criteria essentially linked 
to fluid dynamics and regime analysis, such as gas holdup 
(11)tm ∝
(
Di
N4
)1∕6
,
(12)N2 = N1
(
Di2
Di1
)1∕4
,
parameters and bubble characteristics, liquid properties, 
operating conditions, column dimensions, gas aspersion, and 
characteristics of the solid, liquid, and gaseous components 
of the system [64].
Regarding the design, the BCR is classified as a mul-
tiphase reactor and consists of a vertical vessel where a gas 
or mixture of gases is injected using a nozzle (e.g., spray, 
set of jet nozzles) located at the base of the reactor [66, 
67]. Therefore, aeration and homogenization of the reac-
tion medium are achieved by injecting gas that enters the 
reactor as jets and breaks into bubbles after short distances, 
generating a random movement of the medium that promotes 
gas–liquid mixing [66]. In this context, hydrodynamic prop-
erties and mass transfer are dependent on gas injection and 
sparger flow rates [3]. Airlift reactors, on the other hand, 
are systems derived from BCRs, modified by the presence 
of two channels connected from top to bottom, which allow 
a difference in hydrostatic pressure that induces the circu-
lation of liquid: a channel for upstream gas aspersion (the 
riser) and a channel for downstream circulation of the liq-
uid (the downcomer). This system provides better macro-
scale mixing than a single bubble column. Furthermore, the 
hydrodynamic behavior of this bioreactor configuration will 
be geometry dependent due to the presence of the deflector 
channels for liquid circulation, which can present different 
configurations, being in this context the only controllable 
variable the gas flow rate [3, 65, 66, 68].
Specific phenomena not observable on laboratory scale (due 
to low capacity, low-pressure gradient, and small volumes) 
Fig. 2 Different types of 
pneumatic bioreactors that can 
be applied in bioprocesses. 
Adapted from [55, 65]
372 Systems Microbiology and Biomanufacturing (2024) 4:365–385
1 3
need to be considered when scaling-up. For instance, the pres-
sure gradient along the column increases with increasing liquid 
height, which means there will be a higher pressure along the 
gas nozzle (positive pressure) due to the back pressure being 
applied by the liquid flowing into the nozzle branches through 
the orifices, which will result in a short circuit of gas. Recently, 
the phenomenon was reported by Zhong and collaborators [69] 
who affirm that increasing the scale of the reactor or defining 
a nozzle with a larger number of orifices can improve the bub-
ble distribution and increase the degree of freedom in the gas 
jet. In this scenario, it is worth highlighting one of the critical 
parameters in the scale-up of pneumatic reactors, which are the 
bubble characteristics, since it has a significant impact on the 
hydrodynamics and heat and mass transfers of these bioreac-
tors [64]. It is commonly observed that smaller bubble sizes 
create a larger specific transfer area and are directly related to 
the medium properties, bubble adhesion time, gas rate flow, 
and the diameter of the nozzles used (an essential and critical 
apparatus) [66, 70].
In general, the rising velocity of the bubbles is affected by 
the scale of the reactor, since there is an interaction between 
the vessel walls and liquid characteristics, and the higher the 
liquid column the higher the pressure applied in the gas noz-
zles, which will affect the bubble sizes and dispersion dynam-
ics [3, 69]. The averagebubble diameter (db) can be estimated 
using the equation proposed by Johansen and Boysan [71], 
considering the total gas flow rate (Q) and gravitational accel-
eration (g):
In the scale-up of pneumatic reactors, mass transfer (kLa) is 
one of the most important parameters to estimate and is closely 
related to the gas surface velocity (Usg), which in turn directly 
affects the gas holdup (εG) [3, 66]. Many correlations are pro-
posed to comprehend these dynamics that play critical scale-up 
factors in pneumatic reactors since it is essential to establish 
aeration efficiency and quantify the effects of operational 
variables related to dissolved oxygen delivery [43]. The most 
commonly used correlations to determine kLa in pneumatic 
reactors are presented below, where A, β, and α are dimension-
less parameters, and many β and α are measured by reactor 
dimensions or volumes, type of gas diffuser, and airflow [72].
BCR [73]
Airlift [74, 75]:
The gas holdup is considered a global and dimension-
less parameter that can be related to the dimensions of the 
(13)db=0.35
(
Q2
g
)2
.
(14)kLa = � ⋅ U�
sg
.
(15)kLa = A ⋅ �� ⋅ U�
sg
.
equipment, which can facilitate the design and scale-up and 
can be defined as the volume of the disperse phase (VG) 
divided by the total volume (VL+G), which can also be dis-
criminated by HD as the height of the free surface after aera-
tion (HD), and the height of the free surface before aeration 
(Ho) [3].
Other correlations show that the gas holdup is related to 
the gas surface velocity and the mean bubble rising velocity, 
as shown in the equations below [43].
where VS is the gas surface velocity, and Us is the bubble 
rising velocity.
In airlift reactors, the expression that relates the gas holdup 
in the riser considers the velocity in the core region (VLC) and 
the average linear velocity (VLR). In this equation, the VLC is 
estimated by assumptions about the system, such as assuming 
a parabolic profile for the liquid velocity and the absence of 
gas in the downcomer [43, 76].
In industrial plants, gas–liquid oscillations cause periodic 
changes in the gas flow rate and can improve bubble diffusion 
by increasing the proportion of small bubbles in the reactor 
[69]. Energy requirements are a significant part of the opera-
tional cost of large-scale systems, and it is important to analyze 
the energy input, which in the case of pneumatics is focused 
on the gas injection into the system, and is dependent on the 
global properties of the gaseous and liquid components, and 
can also be associated with the reactor geometry [3, 43, 77, 
78]. In the case of BCRs, if the kinetic energy of the gas flow-
ing out and losses due to attrition are ignored, the following 
equation for pneumatic energy input is accepted [43].
BCR
For airlift reactors, the geometric parameters of the reactor 
are considered and the empirical equation below is accepted.
Airlift
where PG is the energy input by gas injection, VL is the 
liquid volume, ρL is the density of the liquid, AD is the 
(16)�G =
VG
VG+L
=
(HD − Ho)
HD
.
(17)�G =
VS
US
,
(18)�G =
VS
US +
1
2VLC
+ VLR
.
(19)
PG
VL
= �L ⋅ g ⋅ Ug.
(20)
PG
VL
=
�L ⋅ g ⋅ Ug
1 + AD∕AR
,
373Systems Microbiology and Biomanufacturing (2024) 4:365–385 
1 3
cross-sectional area of the downcomer, and AR is the cross-
sectional area of the riser.
In the industrial sector, pneumatic reactors have been 
increasingly used in bioprocesses because they have advan-
tages over conventional reactors (mainly mechanical stir-
ring) by using a single source for stirring and aerating the 
system, providing uniformity and smoothness of turbulence, 
and a simple design and operation with no moving parts 
inside the reaction vessel [67, 79].
Photobioreactors
Microalgae are unicellular or multicellular microorgan-
isms that, unlike most species commonly cultivated in bio-
processes, are highly dependent on light incidence since 
they are photosynthetic. This brings a new element to be 
considered in the design and scale-up of bioreactors—now 
specifically named photobioreactors (PBR) [80]. Photobiore-
actors have some specificity concerning scale-up. Compared 
to STR and pneumatic bioreactors, light is a new variable of 
extreme importance that needs to be considered, and other 
variables must be taken even more rigorously into considera-
tion—such as O2 and CO2 transfers. Oxygen can be toxic to 
microalgae cells above certain concentrations, and CO2 is 
essential for them to perform photosynthesis and also acts in 
the pH regulation of the cultivation media [81, 82]. Different 
configurations of photobioreactors are established today, and 
more are being studied. The most used PBR that have been 
applied for microalgae-based processes are usually vertical 
(bubble column photobioreactor and airlift photobioreactor), 
horizontal (tubular photobioreactor), and flat-panel photo-
bioreactors (Fig. 3).
Horizontal photobioreactors consist of transparent poly-
propylene acrylic, polyvinylchloride (PVC) or low/high-
density polyethylene parallel tubes (10–60 mm of diameter) 
connected to each other [80, 83]. This kind of PBR is suc-
cessfully scaled up to volumes of about 4  m3 or more [80]. 
This type of PBR requires much more power consumption 
than vertical or flat-plate PBR due to the high culture flow 
rate (normally between 20 and 50  ms−1) [80, 83]. Vertical 
column photobioreactors are made of vertical transparent 
glass or acrylic tubes, with a gas sparger at the bottom for 
the effective conversion of the inlet gas to tiny bubbles [80]. 
Generally, the tubes have a diameter of up to 0.1 m to avoid 
limited light availability in the center of the PBR, but in 
scaling-up the diameter can be in the range of 0.2–0.5 m. 
The height of the photobioreactors is constrained to not more 
than 4 m and preferably between 2 and 2.5 m due to struc-
tural engineering motives related to the mechanical strength 
of the construction materials and to prevent mutual shading 
on large-scale cultivations [80, 84].
Amongst the vertical types of PBR, bubble column pho-
tobioreactors offer easy scalability. The sparger design is a 
critical factor in scaling up bubble column photobioreac-
tors since it needs to guarantee microalgae cells protection 
from damage. When using high superficial gas velocities, the 
best strategy for ensuring low gas velocities at the sparger is 
to increase the number of nozzles or increase the diameter 
Fig. 3 Different types of photo-
bioreactors that can be applied 
to microalgae cultivation
374 Systems Microbiology and Biomanufacturing (2024) 4:365–385
1 3
of the nozzles to keep gas velocity lower than the critical 
value [80]. Protective additives can also be added to pre-
vent shear-induced cell damage. The airlift PBR is similar 
to bubble column PBR and has the advantage of preventing 
cell clumping by directing culture media flow in a certain 
direction. This leads to flashing-light effect through the cir-
culation of light and dark zones [80].
One of the most important parameters to consider in 
designing a photobioreactor is the ratio S/V (surface to vol-
ume) since low values of this ratio can lead to insufficient 
microalgae light-harvesting. When compared to horizontal 
surface photobioreactors, vertical photobioreactors have the 
advantage of offering a high S/V ratio, up to 80  m–1 [80], 
which enables reaching higher maximum biomass con-
centrations. This parameter is also related to nutrient and 
gas exchanges in the photobioreactor. A higher S/V ratio 
enhances nutrient exchange, ensuring a more uniform dis-
tribution of essential resources and preventing nutrient limi-
tation and stress zones. It similarly affects gas exchanges, 
as a higher S/V ratio can also positively affect mixing, by 
increasing the contact between the cultivation medium and 
the microalgae. Additionally, the S/V ratio directly impacts 
the scalability and environmental impacts of large-scale cul-
tivationsusing photobioreactors: a higher S/V ratio enables 
reaching increased productivities, which then reduces the 
amount of land and resources required in a large-scale cul-
tivation. In the case of vertical PBR, compact design and 
efficient space utilization is attained [85, 86]. It is impor-
tant to note that while a higher S/V ratio generally offers 
advantages, it is necessary to balance the ratio with other 
design considerations, such as hydrodynamics, mass transfer 
limitations, and practical engineering constraints. From an 
industrial point of view, it is also important to consider that 
the photobioreactor surface area contributes significantly to 
the reactor cost [87].
Flat-panel photobioreactors present a high illuminating 
surface area when compared to horizontal tubular photo-
bioreactors, with a high S/V (surface area to volume) ratio. 
Their modular design is convenient for scaling-up, and the 
agitation of the culture media is provided by either bub-
bling air through a perforated tube or rotating it mechani-
cally through a motor [80, 83]. Flat-panel photobioreactors 
are conceptually designed to make efficient use of sunlight, 
attaining high biomass concentrations, although they occupy 
a considerable superficial area, have complex parts and sup-
port structures, and present difficulties in controlling tem-
perature [83, 88]. In all PBRs, CO2 is normally not only 
furnished with the objective of participating in photosyn-
thesis but also as a way of controlling the media pH. On an 
industrial scale, automatic measures can lead to fresh supply 
of CO2, influencing the amount and form of dissolved carbon 
and bringing the equilibrium back to the ideal conditions for 
each species [81].
Maximizing biomass productivity in horizontal PBRs, as 
well in all PBR types, is strongly related to maximizing the 
irradiance on the surface of the tubes [89]. Molina Grima 
and collaborators [90] proposed that for a fixed biomass con-
centration, the microalgae specific growth rate (μ) depends 
on the average irradiance Iav inside the reactor, according to 
the following equation:
where �max is the maximum specific growth rate, Ik is a 
constant dependent on microalgae species and culture con-
ditions, and n is an empirically established exponent. The 
value of Iav Iav is calculated using the following equation 
[91]:
where Io is the irradiance on the culture surface, �eq is the 
length of the light path from the surface to any point in the 
PBR, Ka is the extinction coefficient of the biomass and Cb 
is the biomass concentration.
For outdoor placed tubular systems, φeq is related to the 
tube diameter φ and the angle of declination (θ) of the sun 
from the vertical [89]; thus,
Besides the importance of the tube's diameter in guaran-
teeing an adequate irradiance to the microalgae, it is clear 
that on an industrial scale the location of the photobioreac-
tors is an essential choice: the average temperature, thermal 
amplitude, and weather of the environment are essential 
for attaining rentable performances in PBR. The geomet-
ric distribution of the tubes also determines the irradiance 
on their surface. Reflectance and shading effects need to 
be accounted for, being the geometric arrangement of the 
tubes an important object of study. An optimal PBR design 
maximizes the amount of solar radiation intercepted and 
distributes it over a larger surface to avoid excess light and, 
thus, photo-inhibition [92].
A key and cost-effective strategy to predict the behav-
ior of PBR parameters in scaling-up is utilizing computa-
tional fluid dynamics (CFD). Those models can help with 
the designing, optimization, and performance evaluation of 
FBR, reducing the dependence on time-consuming and high-
cost experiments [93]. CFD can be used in various aspects. 
In predicting mixing conditions, CFD models can take into 
account transport phenomena and concentration gradients, 
both for nutrients transfer and gases transfer (CO2 and O2), 
which can then suggest the best sparger placement and agita-
tion strategies. Considering hydrodynamics concepts, CFD 
(21)� =
�maxI
n
av
In
k
+ In
av
,
(22)Iav =
Io
�eqkaCb
[
1 − exp (−�eqKaCb
]
,
(23)�eq = �cos�.
375Systems Microbiology and Biomanufacturing (2024) 4:365–385 
1 3
can predict fluid velocity, turbulence, and shear stress, iden-
tifying areas of low flow or stagnant regions that can affect 
the distribution of light, nutrients, and dissolved gases. It can 
also offer insights into the efficient use of light, simulating 
the interaction of incident light with the reactor geometry 
and with the microalgae, considering the optical properties 
of the culture medium, FBR walls properties, photosynthetic 
parameters of each microalgae species, etc. [93, 94].
Besides scaling-up the structure of a PBR, it is important 
to consider the operational mode of it afterwards. In the 
case of a simple batch cultivation, biomass concentration 
and, thus, light attenuation conditions evolve with time. In 
the case of continuous cultivation, biomass concentration 
and light attenuation will directly depend on the dilution 
rate. High biomass concentrations inside the PBR due to 
low residence time can lead to loss of biomass productivity 
and consequently negatively impact the economics of the 
process. It can also cause changes in microalgae compo-
sition and reductions in pigments production, due to high 
light incidence per cell. For that, photon flux density (PFDs) 
larger than 200 μmolm−2  s−1 should be avoided. On the con-
trary, if biomass concentration is maintained too high due to 
a high residence time, dark zones are going to be formed and 
those will hinder biomass productivity [81, 95, 96].
Scale‑up strategies for solid‑state processes
The SSF take place in porous solid supports in which the 
microorganisms grow, and the low water content avail-
able indicates that some of the SmF parameters cannot be 
used for scaling-up, or must be adapted [31, 97, 98]. Each 
SSF scaling-up strategy is rather unique, and the resulting 
industrial-scale bioreactor can have several characteristics 
redesigned. The SSF bioreactors vary from static rectangular 
trays to the vertical columns for packed-bed or fluidized-bed 
bioreactors, and to the horizontal drum and multi-drum bio-
reactors (Fig. 4). Some of SmF parameters associated with 
agitation (when the operation is non-static) and aeration can 
be used as criteria for these vessels. Some SSF are agitated 
by mechanical devices, or even manually, to homogenize 
compounds, heat and oxygen [31, 54]. Yet, the major respon-
sible for heat and mass transfer is the air circulation, directly 
affected by the height of the solid support in the bioreactor 
and its porosity [99, 100]. This indicates parameters associ-
ated with oxygen transfer such as kLa, can be adapted and 
used as scaling criteria [31, 101].
The SSF, by definition, occurs in low water activity rates; 
in other words, the amount of water which is not strongly 
attached to a chemical structure is low. This environment 
characteristic implies that the moisture available must reach 
an equilibrium, as too low water activity can hinder cell 
biomolecules production, transportation and function due 
to denaturation or solute diffusion. The water is present in 
the form of an aqueous film surrounding the microorganisms 
adhered to the solid support, responsible for the transference 
in the microenvironment [97]. However, the major heat and 
oxygen transference occurs by the gaseous phase through the 
pores of the support. Parameters such as flow rate, tempera-
ture and humidity of the air inlet directly affect the system, 
while low porosity and deep height of the bed difficult trans-
fer, especially in the lower layers of the reactor [99]
Before selecting the appropriate values for scaling-up, 
the gradients that are across the vessel should be taken 
into account, as the SSF is not homogeneous. Mathematic 
models that consider oxygen andcarbon dioxide diffusion, 
Fig. 4 Different types of solid-
state fermentation bioreactors a 
tray bioreactor; b rotating drum 
bioreactor; c packed-bed biore-
actor; d fluidized-bed bioreac-
tor; e multi-drum bioreactor
376 Systems Microbiology and Biomanufacturing (2024) 4:365–385
1 3
heat exchanges, distribution of particles, pH, water activ-
ity, substrate consumption, and microorganism growth, or 
even several of those factors, are interesting for selecting 
the best criteria for the larger bioreactor choice [102]Classi-
cal approaches for dimensioning include the trial and error, 
which performs several empirical attempts to verify the 
results, the geometric similarity, maintaining size propor-
tion of the reactors, and the scaling-down method, wherein 
the initial bioreactor to be designed is the industrial and 
then a smaller version is produced. Parameters associated 
with aeration or oxygen transfer such as kLa can be used for 
SSF as well, as the air is constantly percolating the system 
through the solid support pores [31, 101].
The smaller scales for SSF usually occur in glass vessels, 
such as Erlenmeyer and Fernbach flasks, providing a con-
trolled environment for the microorganism to grow. How-
ever, glass bioreactors present a size limit when produced. 
The trays are the most common alternatives for this fermen-
tation, which can be built in plastic, aluminum, bamboo, 
wood or other materials. These trays usually are cultivated 
without agitation in shelves, wherein the height of the bed 
of each container is constant, from 2 to 7 cm. Dimension-
ing this bioreactor type to industry scale is to increase the 
number of trays and shelves, keeping them into a room with 
controlled temperature and moisture, indirectly increasing 
the manual labor for maintenance of all individual fermenta-
tions [54, 103, 104].
Another equipment developed for SSF is the Raimbault 
column, also named as packed-bed or fixed-bed bioreactor. 
In this vessel, the solid support is trapped statically inside 
a column with a nutritive solution layer, all of the system 
over water flasks from where the air is pumped with mois-
ture. Two different strategies utilized to scale up the Raim-
bault column: a dynamic heat transfer model and a modified 
Damköhler number, the former responsible for temperature 
prediction across the bioreactor, and the latter used for calcu-
lating critical size of the bed and different parameters with-
out the necessity of differential equations [31, 99, 105]. The 
Damköhler number considers heat, microorganism growth 
rate, substrate’s density, and can be calculated with the equa-
tion bellow:
where DaM is the Damköhler number (dimensionless), ρS is 
the density of substrate (kg m−3), ε is the void fraction, Y is 
metabolic heat yield coefficient (J (kg dry biomass)−1), µout 
is the specific growth rate at the optimum temperature (s−1), 
Xm is the maximum biomass concentration (kg dry biomass 
(kg initial wet substrate)−1), ρa is the density of moist air (kg 
 m−3), Cpa is the heat capacity of moist air (J kg−1 °C−1), f 
is the rate at which the water-carrying capacity of air varies 
(24)DaM =
0.25�s(1 − �)Y�optXm
�a
(
Cpa + f�
)
VZ(Tout − Tin)∕H
,
with temperature (kg water (kg air)−1 °C−1), λ is the enthalpy 
of vaporization of water (J kg−1), VZ is the superficial veloc-
ity (m s−1), Tout it the temperature of the outlet air (°C), Tin 
is the temperature of inlet air (°C), and H is the bed height 
(m) [105].
Agitated bioreactors are also available for large-scale pro-
duction, including rotating drum bioreactors (RDB), multi-
drum bioreactors, and fluidized-bed bioreactor. The RDB 
consists in a horizontal cylinder using mechanical forces to 
rotate at slow velocities and agitate the culture gently, avoid-
ing hyphae breakage. The drum can reach 200 L of total 
volume, 10 kg of solid substrate, and a common strategy 
to scale up is to insert several cylinders sequentially over 
another, which is called multi-drum bioreactor. It is possible 
to reach 20 kg of solid support by inserting sprinklers over 
the drums, keeping the temperature, moisture and nutrients, 
and the material of the system can be metal or acrylic poly-
mers [54, 103]. Regarding pneumatic agitation, the fluidized 
bed is very similar to Raimbault columns with increased 
forced air to move the solid support. This strategy allows 
better mass and heat transfer, as well as sheer forces, yet do 
not increase solid support capacity [31, 54, 103].
Bioprocess scale‑up examples
Biofuels
The production of many biofuels is established on an indus-
trial scale, with bioreactors scaled up to use different feed-
stocks. Biological processes are conducted in STRs since the 
stirring is responsible for avoiding dead zones in the reac-
tor and increasing the contact between cells and substrates. 
Innovations and scaling-up were increasing over the years 
and intensified after the Paris Agreement (Agenda 2030) 
[106]. Currently, there are several companies on the mar-
ket with large-scale production of biofuels (e.g., ethanol, 
biodiesel, biomethane), being expanded to new sectors that 
have gained great attention in recent years, such as drop-in 
biofuels [107–109].
The wide range of possible feedstocks to be applied 
in biofuel production is a major challenge for large-scale 
development, as they require technological adaptations of 
existing unit operations or the development of new ones. 
In the context of the scaling up of innovative processes 
(either by feedstock or technology), the production of bio-
hydrogen from sugarcane molasses and groundnut deoiled 
cake was carried out in 50–10,000 L scale-up reactors, 
observing that cumulative gas production trends were con-
sistent during process operation [110]. Interestingly, the 
authors highlighted challenges in scaling up, such as clog-
ging of the solid waste recirculation pump, requiring a unit 
operation that results in homogenization and uniformity 
377Systems Microbiology and Biomanufacturing (2024) 4:365–385 
1 3
of the particles; the necessity of installing a moisture trap 
before the gas meter, avoiding blockage of the flow meter; 
possible contamination of inoculum and culture medium, 
which are more challenging in pilot and industrial scales; 
and finally, the high process downtime (maintenance and 
cleaning) after a batch operation, being possible to evalu-
ate processes with a sequential batch operation to avoid 
long downtimes [110].
Another challenge still explored within biofuels is the use 
of residual biomass and the heterogeneity and pretreatment 
of these biomaterials. On a small scale, biomass homog-
enization is generally not problematic, as small amounts of 
samples are mixed and sampled representatively. However, 
this is a challenge on a large scale, discussed and presented 
recently in the study by Adam and collaborators [111], who 
conducted a scale experiment for the efficient blending of 
herbaceous biomasses (leaves and wheat straw) aimed at fuel 
production. The process was conducted with large amounts 
of feedstock, resulting in 28 tons of biomass, which was 
pretreated using a process called florafuel leaching (patent 
WO2009133184) that removes impurities and water from 
the biomass, improving the homogenization that was con-
ducted in a mixer of 2 m3 [111, 112].
Furthermore, in the field of innovation, recently the patent 
WO2019083244 was granted for a method of pretreatment 
and saccharification of biomass to produce biofuels and bio-
plastics, using biological processes of biomass degradation 
before the pretreatment process, denoting a process that, 
according to the authors, requires low economic investment 
and is environmentally sustainable and can reach large scale 
quickly [113]. To solve challenges in biofuel production 
from residual biomass, the granted patent WO2023092956 
proposes a system for cellulosic ethanol production by inter-
mittent saccharification and fermentation of biomass thataims to suppress problems with intermediate products and 
to conduct a process with higher efficiency (also suppressing 
optimal temperature problems of the biological processes 
involved). The systems have coupled reaction and recircula-
tion systems, which allow the control and maintenance of 
temperature, solid loading, and fiber digestion, for efficient 
biofuel production [114].
In biofuels, there is growing interest in the expansion of 
new sectors and the industrialization of new molecules, and 
there is a niche opportunity to develop projects on an indus-
trial scale. Innovations that aim to reduce unit operations and 
maximize product recovery are of great interest to the indus-
try, as shown in patent WO2012079138, where a mechani-
cally stirred reactor and an external jacket for temperature 
control and thermal sterilization of the system were patented, 
coupled with membrane microfiltration modules for passage 
of sterile culture medium and separation of microorganism 
cells after fermentation. This system was patented to obtain 
biosurfactants, biofuels, and enzymes [115].
Also in this context, seeking to solve environmental prob-
lems, the development of biotechnologies aimed at biofuel 
production becomes even more relevant. In 2017 a patent 
(US20170341942) was granted on a large-scale CO2 utiliza-
tion system for utilizing the gases generated by Lake Kivu 
(Africa). This integrated system aims at the production of 
electric energy and storage to produce a range of products, 
among them biofuels. The technology relies on CO2 methane 
degasification and CO2 capture and storage for large-scale 
applications [116]. Although biofuels are already strongly 
impacting the world economy from biobased industrial 
plants, major advances have been observed in the expansion 
of new molecules and technologies, and still concerned with 
the overcoming of problems and challenges still contained 
in large-scale plants. As research advances, it is expected 
that new biofuels will become products of industrial plat-
forms, reducing environmental impacts and moving towards 
a biobased economy.
Food and ingredients
The food industry is one of the areas that SSF is used in 
large scales, together with the enzymatic production. The 
culture of edible mushrooms, such as Agaricus bisporus and 
Lentinula edodes, for instance, is performed in tray biore-
actors, and the industrial scale is reached by enhancing the 
number of trays with the same height of the bed. This bio-
reactor is also used for citric acid and enzymes production 
[54, 103, 104]. The same strategy is applied for the Japanese 
production of natto. The main company responsible for the 
production is the Suzuyo Kogyo Co., and the fermentation 
occurs in individual packages of 50 g for about 18 h after 
the washing and cooking of soy grains. The production size 
only depends on the refrigeration capacity of the industry, 
reaching a production of 238 thousand tons of natto from 
132 thousand tons of soy [117, 118].
Regarding submerged fermentation, the production of 
food can occur in the production of edible mushrooms and 
milk-derived yogurts, for instance. As cultivating mush-
rooms or mycoproteins in SSF takes a considerable time, 
growing them into liquid nutrient solutions provide a higher 
biomass productivity, yet largely modify the final product 
format to propagules. Large production of mycoproteins is 
reported to reach 1,300 L into CSTR with Rushton standard 
impellers, scaling from 75 L bioreactor with geometric simi-
larity [119, 120]. Kefir-derived yogurt, a microorganism’s 
consortium for fermenting milk into beverages, can also be 
scaled up. The consortium biomass production was already 
dimensioned from 1.5 to 2000 L with several steps using 
bubble column bioreactors with conical bottom to collect 
biomass [121]
Several food ingredients and compounds added to 
modify flavor, texture or essence are also produced by 
378 Systems Microbiology and Biomanufacturing (2024) 4:365–385
1 3
microorganisms in submerged fermentation. The erythritol, 
a sweetening agent with 70% of the sucrose’s power, can 
also reach industrial production by scaling-up. The CSTR 
used for its production was increased from 2 to 1500 L with 
four steps using the impeller tip speed criteria and geomet-
ric similarity [122]. Xanthan gum, a polymer that is used 
in food industry for increasing liquid viscosity and stabi-
lizing emulsions, is largely produced by the Xanthomonas 
campestris bacteria. For instance, it was possible to produce 
43.15 g  L−1 of the polymer in 15 L pilot-scale bioreactor 
using fed-batch strategy in 60 h production, almost two-fold 
the batch value of 28.5 g L−1[123]. The xanthan gum can 
also be produced by alternative culture medium, such as 
winery wastewater. The 5 L bioreactor with of 30 g L−1 of 
sugar content in the residue was able to produce of 23.9 g 
 L−1 of the gum in 96 h [124]
The production of enzymes is also frequent in food indus-
try, as they confer different flavor and properties to the final 
product. The β-mannanases, for instance, enzymes utilized 
for fruit beverages and instant coffee, were already scaled 
up into CSTR to a 30 L bioreactor with Rushton standard 
impellers. The fed-batch operating mode was capable of 
reaching 302.6 U mL−1 enzymatic activity from Aspergillus 
sojae in carob pod extract media [125]. Other example is the 
production of amyloglucosidase and exo-polygalacturonase 
by Aspergillus niger. These enzymes can be used to degrade 
gelatinized starch into constituent sugars, and can be pro-
duced in 2 kg rotating drum bioreactors in solid-state fer-
mentation. The process reached 886 U g−1 of amyloglucosi-
dase and 84 U g−1 exo-polygalacturonase from rice bran and 
rice straw [126]The β-galactosidase, also known as lactase, 
is other essential enzyme which hydrolyzes the lactose into 
glucose and galactose, largely applied in dairy products such 
as ice cream and cheese. This enzyme can be produced by 
Bacillus licheniformis in 5 L STR using chemically defined 
medium. The process was able to produce 225.2 U mL−1 
of β-galactosidase in 48 h after optimization, two times the 
production value without optimization [127].
Waste treatment
Bioreactors have gained significant importance in wastewa-
ter treatment since, in addition to removing pollutants, they 
can also generate energy in the form of methane [128], and 
hydrogen [129, 130]. According to Deng et al. [128] the 
main bioreactors used for the treatment of effluents on an 
industrial scale are: completely stirred tank rector (CSTR), 
these systems are widely used in the treatment of domes-
tic sewage, the upflow anaerobic sludge blanket (UASB) 
employed mainly for the anaerobic treatment of industrial 
wastewater. Other types of bioreactors that have also been 
used in wastewater treatment to a lesser extent are upflow 
solids reactors (USR) and upflow blanket filter (UBF) 
reactors.
Extensive works of scaling bioreactors for the treat-
ment of effluents have been developed. For example, [131] 
scaled the methane production process of a bioreactor 
from 0.05 to 500 m3 using palm oil mill effluent (POME) 
as substrate in a bioreactor operated in semi-continuous 
mode using a suspended anaerobic digester. In the proto-
type and in the larger-scale bioreactor the organic loading 
rates (OLR) and methane production rates were: 6.0 kg 
COD m−3  day−1/0.992  m−3  m−3
reactor day−1 and 5.0 kg COD 
 m−3  day−1/∼ 1000 kg biogas/3000 kg COD day−1, respec-
tively; the scale-up criterion was OLR. On the prototype 
and larger scale the COD removal rates were 95 and 97%, 
respectively [131, 132]. The effluent treatment technology in 
bioreactors is a consolidated technology in the industry. In 
China, for example, by the year 2010, around 2842 treatment 
bioreactors had been installed with a capacity to treat around 
130 million m3/day [128]. In Brazil, approximately 5.4 × 105 
 m3/day of wastewater is treated usingUASB; India and Mid-
dle East also use these bioreactors for wastewater treatment 
with a DOC removal efficiency of approximately 80% [133].
Bioactive compounds
There are numerous bioactive compounds that can be pro-
duced using bioprocess production methods (Table  4). 
Some of the main bioactive compounds produced through 
bioprocesses include: antibiotics, enzymes, amino acids, 
and organic acids. Antibiotics are widely used in medicine 
to treat bacterial infections. Stirred tank bioreactors are 
commonly employed, they provide efficient mixing, oxy-
gen transfer, and temperature control, allowing for high 
yields of bioactive compounds [134], as they remain the 
best alternative when the objective is optimizing conditions 
to produce the well-known penicillin and natamycin in fed 
and fed-batch strategies [135]. Solid-state bioreactors have 
potential in antibiotic production, offering novel avenues for 
cost-effective antibiotic manufacturing where the scalability 
is possible using genomic approaches as the amphotericin 
production in a 50-ton bioreactor described by Huag et al. 
[136]. Still, plant cell bioreactors have gained attention for 
the production of complex biopharmaceuticals, as they pre-
sent an alternative for the sustainable production of novel 
antibiotics, emphasizing the controlled environment for plant 
cell growth and secondary metabolite production [137].
Enzymes have a wide range of applications in industries 
such as food, pharmaceuticals, and biotechnology. Exam-
ples of industrially produced enzymes include amylase, 
protease, lipase, and cellulase. For instance, in the produc-
tion of enzymes like α-amylase, cellulase, and lipase, airlift 
bioreactors have shown advantages in terms of mass transfer 
and lower shear stress on the cells or enzymes. Stirred tank 
379Systems Microbiology and Biomanufacturing (2024) 4:365–385 
1 3
bioreactors also remain a great option for enzyme produc-
tion; recent research [138] demonstrated the use of stirred 
tank bioreactors to optimize fructosyltransferase production. 
Advanced agitation and aeration strategies in these biore-
actors have shown promise in improving enzyme yields. 
Solid-state bioreactors were earlier consolidated due to the 
characteristics of this fermentation technique stimulates a 
natural habitat for fungi-enzymes producers, even though it 
offers several advantages, the scalability was always a chal-
lenge. Currently, novel strategies are being designed taking 
into account the crucial parameters such as the generated 
gas distribution and monitoring the variations on the initial 
moisture [139]. Besides, membrane bioreactors (MBRs) 
are used for the production of intracellular bioactive com-
pounds, such as intracellular enzymes or metabolites. MBRs 
combine traditional bioreactor principles with membrane fil-
tration, allowing for the retention of cells or particles while 
the liquid medium is continuously circulated. The strategy 
of continuous enzyme production offers improved product 
quality and recovery efficiency. Packed-bed bioreactors are 
often employed for the production of bioactive compounds 
when using immobilized enzymes or cells. In these bioreac-
tors, the immobilized biocatalysts are packed within a col-
umn or vessel, and the substrate flows through the packed 
bed. Packed-bed bioreactors are used in the production of 
enzymes, biopolymers, and various biochemicals [134].
Amino acids, such as glutamic acid and lysine, are pro-
duced on a large-scale using bioprocesses. They are used as 
food additives, animal feed supplements, and in the produc-
tion of pharmaceuticals and biodegradable polymers. Stirred 
tank bioreactors remain versatile platforms for optimizing 
glutamic acid production, due to their precise control over 
culture conditions, promoting higher yields [140]. MBRs 
bioreactors enable the production of intracellular amino 
acids, particularly when the target amino acids are mainly 
found inside the microbial cells. This type of bioreactor has 
been used for controlling the metabolic flux while amino 
acids are produced, enabling the creation of a metabolomic 
profile to target the best production pathways towards a strat-
egy of anammox process at low temperature, for example, to 
enhance amino acids production [141]. Packed-bed bioreac-
tors are often used due to its configuration provides a high 
surface area for interaction between the immobilized cells or 
enzymes and the substrate, facilitating efficient amino acid 
production [142]. Organic acids include citric acid, lactic 
acid, acetic acid, and malic acid. These acids are used as 
food preservatives, flavor enhancers, pH regulators, and in 
the production of various chemicals and polymers. All types 
of bioreactors mentioned can be employed for producing 
organic acids, including fluidized-bed bioreactors involving 
the suspension of solid particles (e.g., cells or immobilized 
enzymes) in an upward-flowing fluid. The fluid velocity is 
adjusted to keep the particles in a fluidized state, allowing 
for efficient mass transfer and high productivity, they are 
suitable for continuous processes and can be used for organic 
acid production [142, 143]. It is important to note that the 
production of bioactive compounds through bioprocesses 
is a broad field, and there are numerous other compounds 
that can be produced using different microorganisms, plants, 
and bioprocessing techniques. The choice of specific bioac-
tive compounds for production depends on their commer-
cial value, demand, and feasibility of production through 
bioprocesses.
Microalgae bioprocess
Microalgae are highly envisaged for gas mitigation, since 
they consume CO2 for photosynthesis and produce a cleaner 
Table 4 Bioactive compounds 
productions comparing the 
scalability
n.p. not provided
Bioactive compound Production in shake flasks Production in bioreactor References
ε-Poly-l-lysine n.p 2 L
27.07 g/L
[133]
Erythritol n.p 1.5 m3 pilot-scale bioreactor
180.3 kg/m3
[120]
l-asparaginase n.p STR 7 L
162.11 U/mL
[134]
Lactic acid 78.75 g/L 50 L pilot-scale bioreactor
73 g/L
[135]
α-amylase n.p STR 7.5 L
150 U/ml
[136]
Malic acid n.p 7.5 L
95.2 g/L
[137]
l-Arginine Fed-batch 5 L
92.5 g/L
Fed-batch 1500 L
81.2 g/L
[138]
β-Farnesene n.p 300 L pilot-scale bioreactor
900 g/L
[139]
380 Systems Microbiology and Biomanufacturing (2024) 4:365–385
1 3
gas, O2. Since enormous amounts of flue gases are generated 
day by day in industries from burning fossil fuels, large-scale 
PBRs integrated into the combustion processes are aimed. In 
the article by Pereira et al. [144], the scale-up and large-scale 
production of Tetraselmis sp. CTP4 was performed focus-
ing on CO2 sequestration. The authors started with a culture 
in an agar plate, and until reaching the 100  m3 horizontal 
tubular PBR, the microalgae was cultivated for 7 days in 
each of these scales: 100 mL Erlenmeyer flasks, vertical 1 L, 
two 5 L airlifts; 125 L flat-panel, 1 m3 flat-panel; two 2.5 m3 
pilot-scale tubular PBR, industrial-scale 35 m3 tubular PBR; 
and finally 100 m3 tubular PBR.
Optimization was performed in 2.5 m3 tubular PBR: cul-
ture velocities of 0.65, 1.01 and 1.35  ms−1 were tested at 
a fixed pH of 8.0, while three distinct pH set points (7.0, 
7.5 and 8.0) were tested at a culture velocity of 1.01  ms−1. 
The industrial production of microalgae biomass was then 
carried out in 35 and 100 m3 horizontal tubular PBR, with 
a culture velocity of 1.01  ms−1 and a pH set point for CO2 
injection of 8.0. The respective area of implementation of 
the PBRs was 133 and 405 m2, having a total length and 
width of 48.2 × 2.5 m and 96.0 × 4.0 m for the 35 and 100 m3 
PBRs, respectively. The mode of operation was semi-contin-
uous: every 13–14 days approximately 70% of the total cul-
ture volume was harvested while the remaining culture was 
renewed with fresh cultivation medium. The results obtained 
in the two industrial scales were very