Strategies and engineering aspects on the scale-up of bioreactors for different bioprocesses

Abstract

Bioreactors are central equipment used in the majority of bioprocesses. Different models of bioreactors have been developed 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 reproduction 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.

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Vol.:(0123456789)
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Systems Microbiology and Biomanufacturing
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Wedderho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
© 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

Systems Microbiology and Biomanufacturing
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 inuencing 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
Table 1 Advantages and disadvantages of different modes of operation
Operation mode Characteristics Advantages Disadvantages Potential biomolecules References
Batch Fixed volume, without any feeding or
media withdrawal
Easy to control, low maintenance
cost, low chance of contamination
Lower productivity, higher down-
times, and high osmolarity in the
beginning of the process
Bioethanol, organic acids, enzymes,
amino acids
[8, 11, 20–22]
Fed-batch Feeding of carbon source and/or
nutrients, no broth withdrawal
No carbon source inhibition, higher
productivity
Control and feeding strategy deter-
mination can be tricky, end-product
inhibition
Bioplastics (e.g., PHAs), biodiesel,
butanol, bioethanol, among others
[8, 11, 14, 23, 24]
Continuous Both feeding of fresh media and
fermented broth withdrawal, main-
taining equilibrium
Low downtime, high productivity Washout can occur, high chance of
contamination
Organic acids (e.g., succinic),
butanol, bioethanol, biohydrogen
[8, 11, 25–27]

Systems Microbiology and Biomanufacturing
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 aecting 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]

Systems Microbiology and Biomanufacturing
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 oxygen uptake 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

Systems Microbiology and Biomanufacturing
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 DT2/4), 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
𝜋D2
T),
(3)
(kLa)1=(kLa)2.
Fig. 1 STR model with geometric similarity relations. Adapted from [11, 54]

Systems Microbiology and Biomanufacturing
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 growth
rate (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)
N
2=N1
(
Di2
Di1)
2B−2.85A
3.15A
.
(
Q2
Q1)
0.25A−B
3.15A
,
(5)
N
P=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 NDi2). 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)
N
P=
P
N3D5
i
𝜌
,
(7)
N
Re =
ND
2
i𝜌
𝜇
.
(8)
(P
V)1
=
(P
V)2
,
(9)
N
2=N1.
(
Di1
D
i2)2∕3
,
(10)
N
2=N1
(
Di1
D
i2)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

Systems Microbiology and Biomanufacturing
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)
t
m∝
(
Di
N
4
)1∕6
,
(12)
N
2=N1
(
Di2
D
i1)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]

Systems Microbiology and Biomanufacturing
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 average bubble 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)
d
b=0.35
(
Q2
g)2
.
(14)
k
La=𝛼⋅U
𝛽
sg.
(15)
k
La=A⋅𝜀
𝛼
⋅U
𝛽
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)
𝜀
G=
V
G
V
G+L
=
(H
D
−H
o
)
H
D
.
(17)
𝜀
G=
V
S
U
S
,
(18)
𝜀
G=
V
S
US+1
2VLC
+VLR
.
(19)
P
G
VL
=𝜌L⋅g⋅Ug
.
(20)
P
G
VL
=
𝜌
L⋅
g
⋅
U
g
1+AD∕AR
,

Systems Microbiology and Biomanufacturing
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

Systems Microbiology and Biomanufacturing
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-
tivations using 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
𝜇max
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,
𝜑eq
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)
𝜇
=
𝜇
max
In
av
In
k
+In
av
,
(22)
I
av =
Io
𝜑
eq
k
a
C
b
[
1− exp (−𝜑eq KaCb
],
(23)
𝜑eq
=𝜑cos𝜃.

Systems Microbiology and Biomanufacturing
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 and carbon 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

Systems Microbiology and Biomanufacturing
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)
Da
M=
0.25𝜌
s
(1−𝜀)Y𝜇
opt
X
m
𝜌
a(
C
pa
+f𝜆
)
V
Z
(T
out
−T
in
)∕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

Systems Microbiology and Biomanufacturing
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 that
aims 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

Systems Microbiology and Biomanufacturing
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−3reactor 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 using UASB; 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

Systems Microbiology and Biomanufacturing
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]

Systems Microbiology and Biomanufacturing
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 similar. 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
Top t
,
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

Systems Microbiology and Biomanufacturing
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 project administration,
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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