Field of the Invention
The present invention relates to the culturing of microorganisms for
the production of a biosynthetic product. In particular the present invention r ela
tes to a method for producing an intracellular microbial product in a cascade of
bioreactors. Specially, the present invention relates to methods and a cascade of
bioreactor systems comprising at least two bioreactors and at least one concent
ration unit.
Background of the Invention
Microorganisms of different kind can be used for production of many
different kinds of products, ranging from products they will naturally produce to
products they may be manipulated to produce either by genetic engineering or by
selected feed supply. Advantageously, certain microorganisms can be grown on
cheap waste products to produce important valuable products.
US 2012/0219993 (to Chang et al.) relates to microorganisms for
production of intracellular products from volatile fatty acids in multi-stage bior e
actor system, wherein the microorganisms are grown in a growth reactor, and
subsequently culturing the grown microorganisms in a production reactor of biosynthetic
intracellular products. Microorganisms were grown in growth medium,
and subsequently cultured in medium comprising growth inhibitory volatile fatty
acids as a carbon source for production of intracellular products.
Fermentation, or product formation in reactors of the batch type usu
ally suffers the drawback of low productivity due to low cell concentration. Various
attempts have been made to overcome this problem, as well as the problem
of end product inhibition when producing organic acids by microbial fermentati
on or product formation.
US 6,596,521 (to Chang et al.) describes a cell recycle multiple stage
continuous fermentor with serially connected fermentors, wherein the concentration
of microorganisms in a fermentation reactor is increased by filtering off part
of the fermentation media.
There is a need for further methods for producing biosynthetic prod
ucts, either intracellular or extracellular, as well as improving cost effectiveness
in large scale microbial production of biosynthetic products.
Summary of Invention
The present invention was made in view of the prior art described
above, and one of the objects of the present invention is to provide improved pro
cesses for the production of biosynthetic products in a bioreactor system. Another
objective of the present invention is to provide a process, which enables produc
tion of the biosynthetic product with high productivity.
The objects of the invention are achieved by a method which is charac
terized by what is stated in the independent claims. The preferred embodiments
of the invention are disclosed in the dependent claims.
The present invention provides a method for producing a biosynthetic
product, especially an intracellular product, in a cascade of bioreactors, e.g. ser ial
ly connected bioreactors, by utilizing at least one biomass production reactor and
at least one product formation reactor, said product formation reactor being in
flow connection with a concentration device. The method comprises culturing
microbial cells (i.e. cells from one or several microorganisms) under optimal con
ditions for growth, subsequently changing the conditions to allow the cells to en
ter an effective product formation phase. This process of culturing a microor
ganism may according to the present invention be performed in a bioreactor sys
tem comprising an array of at least two bioreactors (e.g. serially connected) so
that cells obtained from culturing said microorganism in one of said bioreactors
are concentrated and partly fed into the same or into the subsequent bioreactor
in the array of bioreactors.
More specifically the present invention relates to a method for produc
ing a biosynthetic product in a cascade of bioreactors (e.g. serially connected b io
reactors), the cascade comprising a bioreactor system for biomass production
comprising at least one biomass production reactor, and a bioreactor system for
product formation comprising at least one product formation reactor in flow con
nection with a concentration device, which method comprises:
a) culturing a microorganism in a biomass production bioreactor by
feeding the bioreactor with a nutrient rich culture medium allowing efficient
growth of biomass;
b) taking at least part of the microorganism culture from the biomass
production reactor of step a) and feeding it to a product formation reactor contai
ning nutrient depleted medium optimized for formation of the biosynthetic pro
duct;
c) producing the biosynthetic product in the presence of the nutrient
depleted product formation medium in the product formation reactor;
wherein the cell mass concentration of the microorganism culture of
the product formation reactor is increased by using a concentration device in flow
connection with the product formation reactor.
The present invention relates also to use of a cascade of bioreactors
(e.g. serially connected bioreactors) comprising at least two or three bioreactors
(e.g. serially connected bioreactors) and one or more concentration devices for
producing a biosynthetic product, wherein one or more concentration devices are
in flow connection with a product formation reactor and arranged between and in
line with two bioreactors.
The inventors of the present invention have found that by concentrat
ing at least part of the cells obtained from the biomass production reactor or the
product formation reactor in a concentration device and by at least partly feeding
the obtained micro-organism rich fraction into the same or into the subsequent
product formation bioreactor in the cascade of bioreactors (e.g. serially connected
bioreactors), the productivity of the biosynthetic production process can be s ig
nificantly improved.
Brief Description of Drawings
Figure 1 presents an embodiment of a culturing process for biosyn
thetic product formation, wherein is presented a cascade of bioreactors with a
first set of parallel bioreactors for biomass production comprising k number of
biomass production reactors, wherein k is between 1 and x, followed by a next
step comprising m number of parallel product formation reactors, wherein m is
between 1 and x. The m product formation reactors in this step are in flow con
nection with or connected to a concentration device for obtaining a microorgan
ism rich fraction of the microorganism culture taken from the product formation
bioreactor. The microorganism rich fraction is recycled to the same bioreactor,
while the microorganism poor fraction is discarded from the process. The first series
of product formation reactors are followed by a next step comprising n num
ber of product formation reactors, wherein n is between 1 and x. Feeds: Feed k : k
= number of the biomass production bioreactor l...x, Feed m: m = number of the
product formation bioreactor in flow connection or connected with a concentrati
on device 1...y, Feed n : n = number of the product formation bioreactor 1...z. Individual
feeds, k (l...x), m (l...y) and n (l...z), can be of the same composition or
some of the feeds are of the same composition or all the feeds are of different
compositions.
Figure 2 presents an embodiment of a culturing process for biosynthetic
product formation, which is similar to that presented in figure 1, but with
the difference that the concentration device is positioned between the biomass
production reactor and the product formation reactor. This has the effect that
while the concentration step still leads to increased cell concentration in the m
number of product formation reactors, the composition of the microorganism cu l
ture in the m number of production reactors will be less influenced by the contents
of the biomass production reactor, as the microorganism poor fraction and
the microorganism rich fraction from these bioreactors is separated before ent e
ring the m number of product formation reactors, whereby only part of the media
from the biomass production reactors will enter the product formation reactors.
Feeds: Feed k : k = number of the biomass production bioreactor l...x, Feed m:m =
number of the product formation bioreactor in flow connection or connected with
a concentration device l...y, Feed n : n = number of the product formation bioreac
tor l...z. Individual feeds, k (l...x), m (l...y) and n (l...z), can be of the same com
position or some of the feeds are of the same composition or all the feeds are of
different compositions.
Figure 3 shows an embodiment of two-step culturing process for biosynthetic
product formation with interim concentration, concentration unit con
figuration 1.
Figure 4 shows an embodiment of two-step culturing process for biosynthetic
product formation with interim concentration, concentration unit configuration
2.
Figure 5 shows an embodiment of cascade culturing process for biosynthetic
product formation with interim concentration.
Figure 6 shows an embodiment of cascade culturing process for biosynthetic
product formation with interim concentration and with varying feeds.
Figure 7 shows an embodiment of cascade culturing process for biosynthetic
product formation with interim cell concentration with fermentors for
nutrient depletion, concentration configuration 1.
Figure 8 shows an embodiment of cascade culturing process for biosynthetic
product formation with interim cell concentration with fermentors for
nutrient depletion, concentration configuration 2.
Figure 9 shows an embodiment of a culturing process for biosynthetic
product formation with multiple interim concentration steps, concentration unit
configuration 1.
Figure 10 shows an embodiment of a culturing process for biosynthetic
product formation with multiple interim concentration steps, concentration
unit configuration 2.
Detailed description of the Invention
In describing the embodiments of the invention specific terminology
will be resorted to for the sake of clarity. However, the invention is not intended
to be limited to the specific terms so selected, and it is understood that each speci
fic term includes all technical equivalents which operate in a similar manner to
accomplish a similar purpose.
Definitions
The term "bioreactor" as used herein refers the reactors in which m i
croorganisms are grown or in which they produce biosynthetic products. Diffe
rent types of bioreactors providing different advantages with regard to opt i
mising culture conditions may be used. In one embodiment of the invention air
lift bioreactors are used in the cascade. According to another embodiment of the
invention both stirred tank bioreactors and air-lift bioreactors may be used in the
cascade. According to another embodiment of the invention only stirred tank b io
reactors are used in the cascade.
As used herein terms "reactor" and "bioreactor" are used int er
changeably.
The term "Product", also referred to as "biosynthetic product", as
used herein refers to biosynthetic products synthesized via natural and/or genet
ically modified metabolism including both intra and extra cellular products that
can be either primary or secondary metabolites.
The term "lipid" refers to a fatty substance, whose molecule generally
contains, as a part, an aliphatic hydrocarbon chain, which dissolves in nonpolar
organic solvents but is poorly soluble in water. Lipids are an essential group of
large molecules in living cells. Lipids are, for example, long chain diacids, hydroxyl
fatty acids, long chain diols, lipids, fats, oils, waxes, farnesene type products, wax
esters, sterols, terpenoids, isoprenoids, carotenoids, polyhydroxyalkanoates, nucleic
acids, fatty acids, fatty acid derivates, fatty alcohols, fatty aldehydes, fatty
acid esters, fatty amines, medium and long chain dicarboxylic acids, epoxy fatty
acids, long chain diols and polyols, phospholipids, glycolipids, sphingolipids and
acylglycerols, such as triacylglycerols, diacylglycerols, or monoacylglycerols. The
terms "lipid", "oil" and "fat" are used in this description synonymously. The term
"acyglycerol "refers to an ester of glycerol and fatty acids. Acylglycerols occur naturally
as fats and fatty oils. Examples of acylglycerols include triacylglycerols
(TAGs, triglycerides), diacylglycerols (diglycerides) and monoacylglycerols (monoglycerides).
The term "microbial lipid" or "microbial oil" refers to a lipid pro
duced, transformed or modified by a microorganism. The term "microbial lipid"
encompasses the category "single cell oil".
In productivity (g/(l*h)) the 1is the bioreactor volume measured in l i
ters and g refers to mass of product in grams.
The term "cell growth curve" refers to which growth state the micro
bial cells are at, e.g. are they in the exponential growth phase or in late exponential
phase or in the stationary phase. The status will indicate which conditions r e
garding nutrition, temperature and aeration are optimal for the cells at the time.
When referring to "optimization of process conditions in individual bioreactors
according to cell growth curve" it is meant that it may be beneficial to amend culturing
conditions in order to shift cells into other conditions, in order to make the
cells enter the next phase, such as going from exponential growth to late exponen
tial growth phase.
The term "cell separation step" or "separation" is a cell concentra
tion step, where part of the microorganism culture is separated from another part
which in consequence will comprise an increased density of cells as compared to
before the separation step. Cell concentration may be achieved by different
means, such as for example filtration, centrifugation, decanting, settling, flocculation
and/or flotation (e.g. as unit operation). Many different types of cell concent
ration devices may be used in the present invention, and the skilled person will be
able to select suitable means.
The term "concentration device" refers to any device, which is used
for obtaining a microorganism rich fraction of the microorganism culture. Op
tionally, in addition to a microorganism rich fraction also microorganism poor
fraction is obtained. Cell concentration may be achieved by different means, such
as for example filtration, centrifugation or by decanting. Many different types of
cell concentration devices may be used in the present invention, and the skilled
person will be able to select suitable means. "A concentration device being in flow
connection with the product formation reactor" refers to a situation, wherein flow
connection is for delivering microorganism culture to said concentration device
or delivering cell mass from said concentration device to the product formation
reactor.
The term "increased cell mass concentration of the microorganism
culture" refers to an increased biomass concentration of the microorganism cu l
ture compared to the microorganism culture before the separation step (as meas
ured in weight % g/1, cfu (colony forming units), or by optical density).
The term "product formation curve" refers to the different phases of
product formation. Cell culturing or fermentation conditions may be different du
ring different phases, why change of media to induce nutrient starvation may be
beneficial, in order to e.g. induce utilization of non-preferred nutrients such as
sugars. Having e.g. certain nutrient or aeration conditions may keep cells in a
product formation mode.
The term "biomass production reactor", refers to a bioreactor in
which temperature, pH, nutrient and aeration conditions are such that cell divi
sion is optimized. The main purpose of a biomass production reactor is to produ
ce biomass. In case of anaerobic growth no aeration of the bioreactor is perfor
med.
The term "product formation reactor", refers to a bioreactor in
which nutrient, temperature, pH and aeration conditions are optimized for prod
uct formation. In case of anaerobic growth no aeration of the bioreactor is per
formed.
The term "aerobic" refers to e.g. cellular culturing under aerobic conditions,
where a microorganism can generate energy in the form of ATP by oxida
tive respiration. The method of the present invention may be run under aerobic,
anaerobic or micro aerobic conditions.
The term "anaerobic" refers to e.g. cellular culturing or fermentation
without oxygen. An anaerobic fermentation process is herein defined as a fermentation
process run in the absence of oxygen or in which substantially no oxygen is
consumed, and wherein organic molecules serve as both electron donor and
electron acceptors.
The term "micro aerobic" refers to oxygen-limited conditions i.e. an
environment in which the concentration of oxygen is less than in air. An oxygenlimited
method is a method in which the oxygen consumption is limited by the
oxygen transfer from the gas to the liquid. The degree of oxygen limitation is determined
by the amount and composition of the ingoing gasflow as well as the a c
tual mixing/mass transfer properties of the fermentation equipment used.
The term "production" refers to the cells metabolic synthesis of m i
crobial products or biosynthetic products, such as microbial transformation of
compounds (e.g. fatty acids to diacids) or by de novo synthesis.
As used herein "intracellular products" refers to any products pro
duced by microorganisms and remaining inside the cell. "Intracellular products"
is used in contrast to "extracellular products".
As used herein "extracellular products" refers to any products produced
by microorganisms and moving or transported outside the cell, e.g. outside
the plasma membrane.
The term "Secondary metabolite" as used herein refers to a product
produced intra or extra cellularly by a micro-organism near the end of the growth
phase or during the stationary growth phase.
The term "Primary metabolite" as used herein, refers to a product
produced intra or extra cellularly by a micro-organism during the growth phase
including metabolites involved in growth, development and reproduction that
typically are energy metabolism related products such as ethanol and lactic acid
are primary metabolites.
The term "Growth" refers to an increase in cell number or an increase
in microbial mass, also referred to as "growth phase".
The term "Exponential growth" as used herein refers to growth of a
micro-organism where the cell number doubles within a fixed time period.
The term "Exponential growth phase" as used herein refers to a period
of the growth cycle of a population in which growth increases at an exponen
tial rate. Also referred to as exponential phase.
The term "Stationary growth phase" or "stationary phase" as used
herein refers to the period during the growth cycle of a microbial population in
which growth ceases meaning no net increase or decrease in the cell number or
microbial mass. In connection to the present invention, stationary growth phase
is also used to refer to the phase during the growth cycle of a microbial populati
on in which slow, but no more exponential, net increase in cell number or micro
bial mass occurs, also referred to as "late exponential growth phase" or "late stage
exponential growth phase" or "early stationary growth phase". Typically, the stationary
growth phase is reached when some growth nutrient becomes growth l i
miting or accumulation of inhibitory metabolic products inhibits cell growth. In
the stationary growth phase, many cell functions may continue, for example ener
gy metabolism and some biosynthetic processes.
The term "Resting cells" as used herein refers to dormant cells, spores
or other microbial resting, non-growing, cell forms. In connection to this invention,
although no growth occurs in resting cells, some metabolic and/or biosynthe
tic processes producing biosynthetic product may occur.
The term "product formation phase" may be defined as a phase, in
which typically at least 80 % of the end product is formed.
The term "media" or "culturing media" refers to the different types of
media used in the present invention. Media or culturing media is suitable for e.g.
biomass production or product formation. In some embodiments of the invention
media for product formation may be the same as the media for biomass producti
on.
The term "microorganism rich fraction" or "cell rich fraction" is the
fraction of the microorganism culture which after separation, i.e. a cell separation
step or a concentration step, comprises an increased biomass concentration com
pared to the microorganism culture before the separation step.
As used herein "microbial" and "microorganism" are used int er
changeably.
The term "culturing" as used herein, refers to the process of growing
the microbial cells under optimal conditions for microbial culture. In one embo
diment of the invention "culturing" refers to any method step of the present in
vention e.g. during the different phases wherein the microbial cells produce b io
mass or the biosynthetic product of the present invention.
The term "culturing conditions" as used herein, refers to all the con
ditions which influence the optimal performance of the microbial cells in either
growth or product formation or in their transition between the different phases
of growth or product formation. As an example, "culturing conditions" include any
aerobic, anaerobic or micro aerobic conditions.
The term "microorganism culture" refers to a culture comprising m i
croorganisms and media for culturing the microorganism cells (e.g. biomass pro
duction or product formation). In some embodiments of the invention the culture
also comprises a biosynthetic product.
The term "fermentation" as used herein, refers to anaerobic fermentation
of microbial cells for biosynthetic product formation.
The term "cascade" or "cascade process" refers to a process of culturing
microbial cells for production of biosynthetic products. The process is per
formed in a set (e.g. cascade) of bioreactors, where the microbial cells may be
moved between bioreactors, in order to allow for changes of the culturing or fermentation
conditions for the cells according to the different steps of the of the
process. The term cascade refers to the flow of cells through the bioreactor sys
tem allowing for a continuous or fed batch culturing and product formation pro
cess.
The term "Single cell oil" as used herein refers to a fatty substance,
whose molecule generally contains, as a part, an aliphatic hydrocarbon chain,
which dissolves in nonpolar organic solvents but is poorly soluble in water. Single
cell oils are an essential group of large molecules in living cells. Single cell oils are,
for example, lipids, fats, waxes, wax esters, sterols, terpenoids, isoprenoids, carotenoids,
polyhydroxyalkanoates, nucleic acids, fatty acids, fatty acid derivates, fatty
alcohols, fatty aldehydes, fatty acid esters, fatty amines, medium and long chain
dicarboxylic acids, epoxy fatty acids, long chain diols and polyols, phospholipids,
glycolipids, sphingolipids and acylglycerols, such as triacylglycerols, diacylglycerols,
or monoacylglycerols. In connection to this invention single cell oil refers to
both intra and extra cellular single cell oil products produced via de novo synthesis
as well as by microbial transformation of culture medium components. In
connection to this invention "lipid" and "microbial oil" are used as synonyms for
single cell oil.
The term "microorganism" encompasses any microorganism capable
of producing a biosynthetic product, e.g. a fungus, such as a filamentous fungus or
yeast, a heterotrophic algae, a bacterium or an archaebacterium. One or several
different microorganisms may be utilized in the present invention.
The microorganism for use in the processes of the present invention
may be any one of:
Filamentous fungi: Penicillium, Aspergillus, Trichoderma, Rhizopus,
Humicola, Zycomycota, Ascomycota, Mucor, Mortierella, Species such as Penicillium
chrysogenum, Trichoderma reesei, Aspergillus niger or Aspergillus terreus
Yeasts: Saccharomyces, Schizosaccharomyces, Candida, Pichia, Rhodosporidium,
Rhodotorula, Cryptococcus, Species such as Candida tropicalis, Sa c
charomyces cerevisiae, Pichia pastoris, Pichia stipilis, Schizosaccharomyces pombe,
Rhodosporidium toruloides or Rhodosporidium fluviale Bacteria: Escherichia, Ba cil
lus, Brevibacterium, Streptomyces, Actinomyces, Arthrobacter, Alcaligenes, Nocardia,
Coryneb acterium, Cupriviadus, Zymomonas, Clostridium, Streptococcus, Rhodococcus,
Ralstonia, Lactobacillus, Species such as Escherichia coli, Bacillus subtilis,
Bacillus licheniformis, Brevibacterium flavum, Corynebacterium glutamicum, Cup ri
viadus necator (Ralstonia eutropha), Rhodococcus opacus, Clostridium ljungdahlii,
Clostridium autoethanogenum, Clostridium butyricum, Clostridium acetobutylicum,
Clostridium beijerinckii.
In non-limiting example, microalgae suitable for the present invention
include Chlorophyceae (recoiling algae), Dinophyceae (dinoflagellates),
Prymnesiophyceae (haptophyte algae), Pavlovophyceae, Chrysophyceae (goldenbrown
algae), Diatomophyceae (diatoms), Eustigmatophyceae, Rhapidophyceae,
Euglenophyceae, Pedinophyceae, Prasinophyceae and Chlorophyceae. More specifi
cally, microalgae may be selected from the group consisting of Dunaliella, Chlorella,
Botryococcus, Haematococcus, Chlamydomas, Crypthecodinium, Isochrysis,
Pleurochrysis, Pavlova, Phaeodactylum, Prototheca, Schizochytrium, Skeletonema,
Chaetoceros, Nitzschia, Nannochloropsis, Tetraselmis and Synechocystis.
In some embodiments, microorganisms capable of producing lipids or
enzymes are used. Such microorganisms are typically a fungus, in particular a fi
lamentous fungus (mold) or a yeast, microalga or bacterium. Lipid producing
molds, dimorphic molds and filamentous fungi comprise, in non-limiting example,
those in the genera Absidia, Aspergillus, Blakeslea, Chaetomium, Cladosporium,
Claviceps, Clodosporidium, Cunninghamella, Emericella, Entomophthora, Fusarium,
Gibberella, Glomus, Humicola, Mucor, Mortierella, Paecilomyces, Penicillium, Puccia,
Pythium, Rhizopus, Saprolegnia, Trichoderma, Umbelopsis, Ustilago and Zygorhynchus,
such as molds of the genus Absidia spinosa, Aspergillus, for example Aficheri,
A. flavus, A. nidulans, A. niger, A. ochraceus, A. oryzae, A. sojae and A. terreus, Bla
keslea trispora, Chaetomium globosum, Cladosporidium herbarum, Claviceps p ur
purea, molds of the genus Cunninghamella, for example C echinulata, C japonica
and C elegans, Entomophthora coronata, Fusarium bulbigenum, Fusarium graminearum,
Fusarium sp., Gibberella fujikuroi, Glomus caledonius, Humicola lanuginosa,
Humicola grisea, molds of the genus Mucor, for example M. circinelloides,
M. plumbeus and M. rouxii, molds of the genus Mortierella, for example M.
isabellina, M. alpina and M. ramanniana, molds of the genus Penicillium, for
example P. javanicum, P. lilacinum, P. spinulosum and P. soppii, Paecilomyces //7acinus,
Puccia coronata, Pythium ultimum, Pythium irregulare, Rhizopus arrhizus,
Rhizopus delemar, Rhizopus oryzae, Ustilago zeae, Ustilago maydis, Zygorhynchus
moelleri, as well as Malbranchea pulchella, Myrothecium sp., Sclerotium bataticola,
Pellicularia practicola, Sphacelothea reiliana, Tyloposporidium ehren bergii, Achyla
americana, Lepista nuda, Tilletia controversa, Cronartium fusiform. Umbelopsis isabellina
and Umbelopsis ramanniana, Umbelopsis vinaceae and Umbelopsis angularis.
Yeasts from the genera, but not limited to Cryptococcus, Trichosporon,
Apiotrichum, Hansenula, Lipomyces, Rhodosporidium, Candida, Yarrowia, Rhodoto
rula, Sporobolomyces, Sporidiobolus, Trichosporon, Torulopsis, Waltomyces, Endomyces,
Galactomyces, Pichia or Cryptococcus, such as Cryptococcus curvatus, Cry p
tococcus albidus, C. terricolus, Trichosproron cutaneum, Lipomyces starkeyi, L. // o
fera, Rhodosporidium toruloides, Candida curvata, Yarrowia lipolytica, Rhodotorula
glutinis, Rhodosporidium fluviale, Rhodosporidium diobovatum, Rhodosporidium
kratochvilovae, Sporobolomyces, Candida, Candida sp. 107, Lipomyces sp. 33, Rho
dotorula gracilis, Trichosporon pullulans or Tfermentans can also be used.
Producing a biosynthetic product
This invention describes a novel, cost efficient cultivation process for
production of biosynthetic products, especially intracellular products, using cu l
tured microorganisms.
The microorganisms may produce the biosynthetic product with or
without being induced to do so, in example via induction by environmental condi
tions, such as presence or absence of certain nutritional factors, such as carbon
source or vitamins or other factors in the media, or by changes in temperature,
pH, dissolved oxygen or other. Additionally, in some embodiments the microor
ganisms may be genetically modified to be able to produce the biosynthetic product,
in example by introduction of a recombinant gene capable of encoding the d e
sired biosynthetic product, or wherein the recombinant gene expresses an en
zyme which is capable of catalysing the synthesis of the desired biosynthetic
product. Recombinant DNA techniques for creating genetically modified microor
ganisms are well known in the art, such as in Sambrook and Russel (2001 "Molecular
Cloning: A laboratory Manual" (3rd edition), Cold Spring Harbor Laboratory
Press.
The cultivation process is based on the use of a bioreactor cascade
comprising at least two bioreactors (i.e. the biomass production bioreactor and
the product formation bioreactor) connected in series and one or more cell concentration
devices arranged between and in line with two serially connected bioreactors.
The cell concentration device may also be arranged after the last b io
reactor, or in connection to a bioreactor.
The bioreactor cascade comprising of at least two serial connected bioreactors
may specifically be operated as a cascade process. The cascade process
is easily adaptable when different types of feeds are needed in different process
phases, such as growth and production phases. Different feeds are required e.g.
the cultivation of biomass requires nutrient rich feed and the product formation
phase requires a nutrient deficient feed for example to achieve nutrient starva
tion. Product formation may also be induced by other changes in cultivation conditions,
such as pH, temperature, aeration conditions or addition of an inducing
chemical. Cultivation of biomass may also be sugar dependent while the product
formation phase may require a different kind of feed, e.g. oils and fats in the feed.
The production of a biosynthetic product can generally be presented
to comprise a two phases: a growth phase in which the amount of biomass increases
and a production phase, in which major part of the biosynthetic product is
generated. The growth phase may be conducted in one or several bioreactors
(growth reactor(s)) and the production phase can be conducted in one or several
bioreactors. A culturing medium may be fed into the biomass production bior eac
tor and medium for product formation may be fed to the product formation reactor.
Alternatively, a culturing medium or medium for product formation may be
fed also to additional bioreactors.
The present invention provides a method for producing a biosynthetic
product in a cascade of bioreactors comprising at least two bioreactors in flow (or
fluid) communication with each other, which method comprises:
providing in the biomass production bioreactor a culturing mixture
containing culturing medium and cultured microorganism;
providing in the product formation reactor (arranged after the b io
mass production bioreactor) a mixture containing medium for product formation,
and microorganisms;
feeding at least part of the culturing mixture to the bioreactor, wherein
the concentration of cells in the bioreactor is increased using a concentration d e
vice arranged before or in connection to the bioreactor.
In some embodiments the method further comprises taking part of the
culturing mixture from the biomass production bioreactor, concentrating the culturing
mixture taken from the biomass production bioreactor to obtain a liquid
fraction and a microorganism rich fraction and feeding the microorganism rich
fraction into the subsequent bioreactor.
In some embodiments the method further comprises taking part of the
mixture from the bioreactor for biosynthetic product formation, concentrating
the mixture taken from said bioreactor to obtain a liquid fraction and a microor
ganism rich fraction and feeding the microorganism rich fraction into the bioreac
tor or to a subsequent bioreactor in the bioreactor system.
Each bioreactor used in the present invention may be a simple con
tainer or a tank (e.g. an air lift type or stirred container or tank). The bioreactor
may be used both as a growth reactor for growth of biomass as well as a product
formation reactor for production of biosynthetic product. In anaerobic fermenta
tion no aeration of the bioreactor is performed. It is preferable to be able to cont
rol one or more parameters such as temperature, pH, aeration (in case of aerobic
reactor), agitation, pressure, flow rate, dilution rate, stirring, inlet, outlet redox
potential, cell density and nutrient contents, as the state of the cells such as whet
her the cells will be growing or producing may depend on these factors. Sp ecifi
cally, in production of the biosynthetic product, the step of shifting from the m i
crobial cell culture from the growth phase to the production phase occurs
through inductive operation by changing one or more of the above factors or by
adding an inducing factor such as a chemical. The products can be formed by pr i
mary or secondary metabolism. In a specific embodiment, the microbial growth
phase and product formation phase are conducted in separate bioreactors, to a l
low for convenient and accurate control of culturing and product formation con
ditions, as well as for having a continuous production system. Further, biosynthetic
products may be produced as intracellular or extracellular products, typi
cally intracellular products are produced.
The product or products of production phase may be gaseous. The ad
vantage of the present innovation is that by using high cell densities in production
phase the product concentration in the product gas is increased.
The microorganisms are cultured and/or fermented in an array of at
least two bioreactors. These at least two bioreactors are in fluid communication
with each other, for example they may be connected with appropriate conduits
having valves, as well as being connected through other intermediary means b et
ween the bioreactors, such as a separation stage. In one embodiment of the invention
each bioreactor is in fluid connection only with the subsequent and previous
bioreactor in the cascade, the bioreactor system for product formation being arranged
downstream of the bioreactor system for biomass production. Having a
cascade of bioreactors, between which cells may continuously be transferred, a l
lows for continuous cell production, product formation and harvesting under op
timal conditions for the cells both during growth and during product formation.
In some embodiments, the bioreactors where the product formation steps are
performed, are operated as any one of continuous, fed-batch or a combination of
continuous and fed-batch. In one embodiment, microorganism cells are produced
by fed-batch mode in one bioreactor, microorganism rich fraction is obtained and
furthermore, said microorganism rich fraction is utilized for producing the biosynthetic
product in another bioreactor.
In some embodiments the array of bioreactors comprises three or four
or more bioreactors, which may provide certain advantages. With reference to fi
gure 6 which presents an array of five bioreactors, which may be optimized for
production of secondary metabolites. In figure 6, two biomass production reactors
are used, followed by a concentration step after which cell concentrate (i.e. a
microorganism rich fraction) is lead into the first of three product formation r eac
tors. This type of setup allows the first biomass production reactor to be opt i
mized for exponential growth, from which cells and media (i.e. a culturing mixtu
re) is transferred into the second biomass production reactor, wherein a second
type of growth media (Feed 2) is added allowing for initial nutrition depletion,
and thereby inducing cells to change to product formation phase. After the second
reactor, cells are concentrated in a concentration step, which produces medium
permeate (microorganism poor fraction) and a cell concentrate (microorganism
rich fraction). The cell concentrate is fed into the third bioreactor which is the
first product formation reactor, in which a nutrient depletion media is added to
the cells. The medium permeate from the cell concentration step may be recycled
or discarded. From the third bioreactor the cells flow into the fourth bioreactor
and from there into the fifth bioreactor after which they may be transferred into
further reactors or harvested for extraction of products. In each further bioreactor,
media with different nutrient content may be present, to allow for optimizati
on of nutrient utilization and product formation. In example, cells may be carbon
starved, leading to utilization of non-preferred sugars. In some embodiments, the
biosynthetic product is an extracellular product, which may be extracted from the
medium permeate.
Accordingly, the present method comprises one or more of the follow
ing steps:
STEP A
Step a) of culturing a microorganism in the presence of culturing medium
in a biomass production reactor.
The biomass production reactor may be dedicated to induction of ex
ponential growth of the particular microorganism. When the bioreactor is opera
ted as a growth reactor, it may be done by feeding a culturing medium comprising
the optimal amount of nutrients. This may be achieved by continuously feeding
the bioreactor with new culturing medium, and continuously allowing cells to
flow to the next bioreactor.
During growth, the microorganism will be using the culturing medium
by metabolising the nutrients present in the medium, whereby the first bioreactor
contains a microorganism culture (i.e. culturing mixture). The culturing mixture
comprises the medium as well as the cultured microorganism. The advantage of
this step in a continuous bioreactor set up, is that cell growth may be kept at a
constant high level, and that cells may continuously be fed to the next bioreactor
whether it is a further biomass production reactor or a product formation reactor.
Microorganism cells may also be produced by fed-batch mode in the biomass
production reactor.
Step a) may be followed by feeding the microorganism culture flow
from the biomass production reactor to a concentration device (i.e. concentrating
the microbial cells). Indeed, at least part of the microorganism culture is taken
from the biomass production reactor to be fed to the concentration device. Taking
of microorganism culture from the biomass production reactor may in example
be done in response to either high cell density or nutrient depletion in the b io
mass production reactor. In a continuous bioreactor system, cell density and
nutrient content will be surveilled, and in order to keep such factors at a level
which will keep cells growing exponentially, the speed by which culture mixture
is removed from the biomass production reactor may be regulated accordingly. If
cell density is too high, culture mixture may be removed and new media added.
The step of feeding the microorganism culture flow from the biomass
production reactor to a concentration device thus includes taking microorganism
culture (i.e. mixture of the biomass production reactor) from the biomass production
reactor, and separating the cultured microorganism from the microorganism
culture taken from the biomass production reactor, and thereby obtaining a microorganism
rich fraction of the microorganism culture, and a microorganism
poor fraction of the microorganism culture.
STEPS Band C
Step b relates to feeding, which means transferring an amount of the
microorganism culture from the biomass production reactor of step a), or an
amount of the microorganism rich fraction of the microorganism culture, and a
medium optimized for product formation to the product formation reactor. Step
c relates to production of the biosynthetic product in the presence of the product
formation medium in the product formation reactor. The microorganism cells
using the medium for product formation, in the product formation phase, only
grow very little, and use the nutrients in the medium for production of biosynthe
tic products. An advantage of feeding the microorganism rich fraction obtained
from the microorganism culture of the biomass production reactor to the product
formation reactor, is that certain nutrients often need to be in low amount in the
medium for product formation, and by removing part of the medium in the con
centration/separation step, a medium comprising low amounts of certain
nutrients may be added as the medium for product formation to adjust the con
centrations of e.g. nitrogen in the medium for product formation, and thereby
stimulate the biosynthetic product formation.
Mixing the microorganism rich fraction with product formation media
is thus done, whereby the product formation bioreactor contains a microorgan
ism culture (i.e. mixture of the product formation bioreactor) comprising the me
dium optimized for product formation, together with the microorganism and the
biosynthetic product which may be either intracellular or extracellular.
One embodiment of the invention comprises a step of taking part of
the mixture from the product formation reactor; and separating the microorgan
isms from this mixture, thereby obtaining a microorganism rich fraction and a m i
croorganism poor fraction.
One embodiment of the invention comprises a step of recycling the microorganism
rich fraction to the product formation reactor. Recycling or feeding
of the microorganism rich fraction to the product formation reactor may allow for
further nutrient depletion of the medium for product formation, and thereby in
duce utilization of less preferred nutrients, e.g. sugars, resulting in a more comp
lete use of the medium for product formation. Further, the advantage of this is
that productivity of the microorganism may increase, as illustrated in the examp
les.
In a specific embodiment, a concentration/separation step is arranged
between or in line with a growth reactor and a product formation reactor. It has
been found that the productivity of the process is significantly improved if the
cells obtained from a growth reactor are concentrated before feeding them to a
product formation reactor or if the cells obtained from a product formation r eac
tor are concentrated before recycling them back to said product formation r eac
tor or feeding them to the subsequent product formation reactor thereby incr ea
sing the cell concentration in the product formation reactor. The effect is especial
ly significant, if the biosynthetic product is produced through a secondary metabolite
reaction route, in which the shift from growth to production phase has to
be induced, e.g. by changing process conditions. The microorganisms may be in
duced to shift from growth to production phase by changing the process condit i
ons, e.g. by changing the content of one or more nutrients in the medium for pro
duct formation by (nutrient starvation), or by other means such as by addition of
an inducer, or by reduction of bioreactor volume.
The microorganism rich fraction of the microorganism culture or culturing
mixture obtained from concentration steps, may have an increase in biomass
concentration as compared to the biomass concentration in the bioreactor
from which the mixture originated (as measured in weight %, g/1, cfu (colony
forming units), or by optical density) of at least 50%, more specifically of at least
100%, such as any one of 10%, 20%, 30%, 40%, 50 %, 60%, 70%, 80%, 90%, or
more specifically by at least 100%.
An example of one setup of the present invention is shown in figure 4,
where an amount of the culturing mixture of the biomass production reactor is
fed to the product formation reactor (denoted product formation, nutrient deple
tion, concentration in figure 4), and where a medium for product formation may
or may not be added (denoted Feed 2 in figure 4). Cell concentrate of the product
formation reactor is recycled back to the same product formation reactor. The
bioreactor set up depicted in figure 4 is particularly well suited for intracellular
products, where the concentration step may help remove inhibitory compounds,
and adjust media composition or microorganism culture and cell density in the
bioreactor during the product formation phase.
Another example of the setup of the present invention is shown in fig
ure 3, where an amount of the microorganism rich fraction of the biomass production
reactor is fed to the product formation reactor (denoted product for
mation, nutrient depletion in figure 3), and where a medium for product formation
may or may not be added (denoted Feed 2 in figure 4. Here, a separa
tion/concentration step between the biomass production phase and the product
formation phase or after the product formation phase allows exchange of medium
or microorganism culture from the growth optimised medium or microorganism
culture to a production optimised medium or microorganism culture. This may be
advantageous, e.g. if the growth medium or microorganism culture comprises me
tabolites or nutrients which may be inhibitory for product formation, or which
might influence the growth state of the cell/microorganism, and thereby inhibit
the induction of product formation. Such inhibitory effects of metabolites or
nutrients are well known in the art. The method as shown in Figure 3 is especially
suitable for production of extracellular products.
In some embodiments of the invention, at least part of the microorgan
ism culture is taken from the biomass production reactor and fed into the concen
tration device arranged between the biomass production reactor and the product
formation reactor to obtain a microorganism rich fraction, which is fed into the
product formation reactor, and at least part of the microorganism culture from
the product formation reactor is taken and fed into the concentration device ar
ranged downstream of the product formation reactor to obtain a microorganism
rich fraction, which is fed into a subsequent product formation reactor (down
stream) in the cascade. One advantage may be an efficient change of culturing/
medium for product formation between bioreactors, e.g. to remove inhibito
ry compounds and maintenance of cell density. This setup also allows for an effi
cient switch from growth to production phase of the cells.
In some embodiments of the present invention an amount of the m i
croorganism rich fraction is recycled to the biomass production reactor or to the
product formation reactor, such as for example between 0% and 50%, such as at
least 40%, or at least 30% or at least 20% or at least 10% is recycled to the b io
mass production or product formation reactor, while 50% to 100%, such as all
cells that are not recycled to the biomass production or product formation r eac
tor, respectively, is carried forward to the next bioreactor.
In some embodiments, if nutrient content is not optimal, culture medium
or microorganism culture of the product formation bioreactor may be r e
moved, and new media may be added to change the content of the relevant nutr i
ents.
In some embodiments of the present invention a third bioreactor con
nected serially to the product formation reactor is provided. Expanding the array
with a third bioreactor allows for additional production and versatility of the
methods of the present invention. Aconcentration step may be present between
the product formation reactor and the third bioreactor, for example as illustrated
in figure 10, but a concentration step between the product formation reactor and
third bioreactor may equally also be absent. The presence of a separation step
may allow increasing the microorganism/cell concentration in production r eac
tors thus enabling a faster or more efficient utilization of carbon sources, which
improves the productivity of the production process. In processes, where the sp e
cific productivity (g/(g*h)) of the biosynthetic product, especially intracellular
product, is low or the microorganism is slow growing, increasing the concentrati
on of cells in the production phase enables higher productivity.
In some embodiments of the present invention, when the series of bioreactors
comprises from three, four or more bioreactors, there will only be a s in
gle separation step (concentration step) between the last biomass production r e
actor and the first product formation reactor, or in connection to the first product
formation reactor, such as exemplified in figures 5, 6 and 7. In some embodi
ments, where an array of three, four or more bioreactors are used, an additional
separation/concentration step after the third, fourth, or any subsequent bioreac
tors may be present, feeding the cell rich fraction back into bioreactor 3, such as is
shown in figure 10.
In some embodiments of the present invention the method additional
ly comprises: i) taking part of the mixture from the product formation reactor; ii)
separating the microorganism from the mixture taken from the product for
mation reactor, thereby obtaining a microorganism rich fraction of the mixture
and a microorganism poor fraction of the mixture; iii) feeding an amount of the
microorganism rich fraction of the previous step optionally together with a s ec
ond medium for product formation or together with the same type of media to the
third bioreactor containing the microorganism, the microorganism using the me
dium for product formation, whereby the third bioreactor contains a microorgan
ism culture (i.e. mixture) of the third bioreactor, comprising the medium, the m i
croorganism and the biosynthetic product. This separation/concentration pro
cedure allows for a change of culturing conditions or for supplementing with certain
nutrients in the late production phase, whereby cells may be switched to metabolize
for example a different carbon source than that which is used in the pr e
vious bioreactor.
In some embodiments of the present invention the cascade of bioreactors
additionally comprises a fourth bioreactor connected serially to the third bioreactor.
One example is provided in figure 5, with four serially connected bioreactors,
of which the first is the biomass production reactor, and the three bioreactors
are product formation reactors. A separation/concentration step may be
present between any of the bioreactors, as well as after the fourth bioreactor. The
concentration step may be present between the product formation reactor and
third bioreactor, and either recycles the cell concentrate (i.e. the microorganism
rich fraction) in full or in part to the product formation reactor, or forwards the
cell concentrate in full or in part to the third bioreactor. This bioreactor set up a l
lows for a production method with four steps in which the culturing or product
formation conditions may be varied in each step. This includes, for example, a
biomass production step in the biomass production reactor with conditions for
exponential growth, a product formation step in the product formation reactor
with up-concentration of cells and nutrient depletion, a third and fourth bior eac
tor for product formation. In this latter process, conditions may be as shown in fi
gure 9, wherein the same feed is used in bioreactors 2-4, or it may be a process in
which different feeds are added to bioreactors 2, 3 and 4. The latter situation a l
lows for change of conditions, to increase productivity, and utilization of different
carbon sources.
In some embodiments of the present invention the method additional
ly comprises: g) taking part of the mixture from the third bioreactor; h ) separating
the microorganism from the mixture taken from the third bioreactor, thereby
obtaining a microorganism rich fraction of the mixture and a microorganism poor
fraction of the mixture; i) feeding an amount of the microorganism rich fraction,
optionally together with a (third) medium for product formation to the fourth b i
oreactor containing the microorganism, the microorganism using the medium for
product formation, whereby the fourth bioreactor contains a microorganism cu l
ture (i.e. mixture) of the fourth bioreactor comprising the medium, the microor
ganism and the biosynthetic product.
In some embodiments of the present invention additional bioreactors
are provided, where each additional bioreactor is connected serially to the preceding
bioreactor, and where the number of additional bioreactors are, 1, 2, 3, 4, 5
or 6. Cell separation or concentration steps may be present between any two bioreactors,
and may feed the cell rich fraction into either the previous bioreactor or
into the next bioreactor, as described herein with reference to the third and
fourth bioreactors.
In some embodiments of the present invention the separation e.g. in a
concentration device arranged between the biomass production reactor and the
product formation reactor and/or in a concentration device arranged between
the product formation reactor and the third bioreactor and/or in step h is con
ducted using means selected from the list consisting of: centrifugation, filtration,
settling, flocculation, flotation, decanting. In a specific embodiment, the cultured
cells are concentrated by means of any one of filtration or centrifugation or by d e
canting. In a specific embodiment, the cell concentration is increased in the con
centration step by at least 10%, 20%, 30%, 40%, 50%, 60%, 70%, 80%, 90%, mo
r e specifically by at least 100% when compared to the concentration in the b io
reactor from where it came. The inventors have found that the cells may be concentrated
by at least 10% in order to have a process with improved productivity
of the biosynthetic product, especially intracellular product. A microorganism
rich fraction obtained from the concentration device arranged as described in this
paragraph may be recycled to the same bioreactor where the microorganism cu l
ture was taken from or fed to the subsequent bioreactor.
In some embodiments of the present invention the separation is oper
ated by state of the art, which may include continuously, semicontinuously or
batch wise operated concentration devices. This may for example be achieved by
using any one of the above suggested separation means, such as centrifugation,
filtration, decanting, settling, flocculation or flotation. Any suitable centrifugation,
filtration (e.g. depth filtration or membrane filtration such as microfiltration, u lt
rafiltration, nanofiltration, diafiltration, reverse osmosis including both dead-end
and crossflow type arrangements), decanting, settling, flocculation or flotation
methods and means are well known to a person skilled in the art and may be ut i
lized in the present invention.
The entire bioreactor system may also be operated in a continuous
manner, and in addition to the continuous operation of the separation step(s) all
the bioreactors are also run in a continuous mode. In addition to means for run
ning the separation step(s) in a continuous manner, means allowing for a cont i
nuous flow of cells between bioreactors have to be employed as well. Alternatively,
the method of the present invention may be operated in a fed-batch manner,
wherein the concentration may be operated in a batch or continuous manner d e
pending on the concentration technique used.
In some embodiments of the present invention an amount of the m i
croorganism culture in each of the serially connected bioreactors is fed into the
subsequent bioreactor of the array of serially connected bioreactors. This pro
cedure allows for a flow of microorganisms all the way through the bioreactor
system, to expose the microorganism to all the different conditions in the diffe
rent bioreactors, whereby the organism is led to produce the biosynthetic pro
duct, especially intracellular product, efficiently. It also allows for an amount of
fresh microorganism obtained in the culturing step to be successively carried
through all the bioreactors.
In some embodiments of the present invention particular reaction
conditions are provided to the bioreactors. For example, independently of each
other the biomass production bioreactor, the product formation reactor, as well
as the third and fourth bioreactors (if present) may be optimised according to the
following. A bioreactor preceding the biomass production reactor, and serially
connected to or in flow connection with the biomass production reactor, may ad
ditionally be present, and may be optimised for microorganism growth. The b io
mass production reactor may be optimised for microorganism growth at late stage
exponential phase and/or optimised for induction of biosynthetic product
formation e.g. by nutritional starvation. The product formation reactor may be
optimised for biosynthetic product formation at a stationary growth phase, op
tionally under nutrient starvation conditions. The third bioreactor may be opt i
mised for biosynthetic product formation with minimal carbon source addition.
In some embodiments the method comprise the following steps: j) b i
oreactor preceding the biomass production reactor and serially connected to or in
flow connection with the biomass production reactor, where the bioreactor pr e
ceding the biomass production reactor is optimised for microorganism growth; k)
The biomass production reactor optimised for microorganism growth at late
stage exponential phase and/or optimised for induction of nutritional starvation;
1) the product formation reactor optimised for biosynthetic product formation at
a stationary growth phase, optionally under nutrient starvation conditions; m)
the third bioreactor optimised for biosynthetic product formation with minimal
carbon source addition, which results in residual sugar consumption. This procedure
may in one embodiment be in a fed batch culture, where cell separation
steps ensure that media, microorganism culture and culture conditions may be
changed between each step. The procedure is performed in a cascade bioreactor
system, such as a fed-batch or continuous culture bioreactor setup, where steps j)
to m are performed in a cascade.
In some embodiments only two or three of the: preceding (j), biomass
production (k , product formation (1) and third (m bioreactors are present.
In some embodiments, the method comprises the following steps: j),
k , 1), and m , with a cell separation step between steps k and 1) wherein the cell
rich fraction is fed into 1), for example as illustrated in the array shown in figure 7.
In some embodiments, the method comprises the following steps: j), 1)
and m , with cell separation steps positioned after each of 1) and m , and wherein
the cell rich fraction is fed back into each of 1) and m , for example as illustrated in
the array shown in figure 10.
In some embodiments of the present invention the cell mass concen
tration is increased by one or more method steps or one or more concentration
devices, wherein the microorganism rich fraction is obtained at a temperature b e
tween 3 to 80°C (e.g. 3 to 30°C or 31 to 80°C) lower than the bioreactor from
which the microorganism culture was taken from and/or wherein the tempera
ture of the microorganism culture to be fed into the concentration device is low
ered by 3 to 80°C (e.g. 3 to 30°C or 31 to 80°C).
In some embodiments of the present invention the one or more sepa
rations into a microorganism rich fraction and a microorganism poor fraction is
conducted at a temperature in the range between 1 to 30°C lower (e.g. 3 to 30°C
lower) than the bioreactor from which the mixture to be separated was taken
from and/or where the temperature of the mixture to be separated is lowered by
1 to 30°C (e.g. lowered by 3 to 30°C). In a specific embodiment, the temperature
of the microorganism culture obtained from a bioreactor is lowered before
and/or during a concentration step. Without wanting to be bound by any particu
lar theory, the inventors believe that the concentration step may cause stress to
cells thereby reducing productivity in the later steps and that a temperature decrease
of the microorganism/cells alleviate the stress to a certain extent compared
to no temperature decrease. This maintains the activity of the cells, which impro
ves production efficiency of cells in the bioreactors.
In one specific embodiment, the temperature of the microorganism
culture is lowered by at least 1°C, such as between 1 and 5°C, or 1 and 10°C, or 1
and 15°C or 1 and 20°C or between 1 and 30°C, such as between 5 and 15°C, or
between 15 and 25°C before or during the concentration step.
In another specific embodiment, the temperature is lowered by at least
1%, such as any one of at least 2%, 3%, 4%, 5%, 10%, 15%, 20%, 25%, 30%, 40%,
50%, 60%, 70%, 80%, or at least 90% of the temperature in the bioreactor from
which the cells were derived, as measured in °C. In a specific embodiment, the
temperature is normalized after the concentration step, i.e. returned to the or igi
nal temperature or changed to a specific temperature optimized for production
formation, wherein the temperature is typically between 20°C and 80°C, specifi
cally between 25°C and 37°C. Some advantages associated with a temperature r e
duction step may be considerably lower cell loss during the concentration step.
Furthermore, the temperature reduction step may in addition shift the microor
ganism/cells into a more stationary phase, thereby promoting biosynthetic pro
duct formation.
Furthermore, in some embodiments of the invention in case the m i
croorganism is a transgenic organism and the product being a product of the
transgene which is controlled by a heat sensitive promoter, such as a heat shock
promoter, a temperature shift (i.e. increase of temperature) may be used to in
duce expression of the product. In such cases, a temperature changes during
and/or after the concentration step may be used to induce expression of the pro
duct. Different heat shock promoters may require different temperature shift levels
to be activated. In example, in Zebrafish, a 9.5°C from 28.5°C to 38°C raise in
temperature induce gene expression greatly in heat shock regulated genes. In
Drosophila cells, the optimal temperature shift is from e.g. 25°C to between 36
and 37°C, whereas in Saccaromyces cerevisiae the optimal temperature shift is
from e.g. 25°C to between 39 and 40°C. So in some embodiments, the microorganism
used in the present invention is a transgenic or transfected microor
ganism wherein the transgene or transfected DNA comprises a gene of interest
for production of a biosynthetic product, and wherein the expression of the gene
of interest is controlled by a heat shock promoter. In such embodiments where
the gene of interest is controlled by a heat shock promoter, the temperature before
or during the production phase may be raised by at least 2°C, such as at least
3°C, or at least 4°C, such as at least 5°C, such as between 2 and 15 °C, or between 5
and 15 °C, such as any one of 5°C, 6°C, 7°C, 8°C, 9°C, 10°C, 11°C, 12°C, 13°C, 14°C
or 15°C. The temperature increase may be temporary to induce expression, or
may be permanent during the production phase.
In some embodiments, expression of an introduced gene of interest is
controlled by a carbon source dependent promoter, such as in Saccaromyces cerevisiae,
where certain promoters may be induced by lack of glucose (Weinhandl et
al. Microbial Cell Factories, 2014, 13:5.). Other useful promoters which may be
induced by a variety of factors are well known in the art, in non-limiting example,
such a promoter may be the yeast Tet-Promoters which are inducible by addition
of tetracycline may be used. Copper ion inducible promoter systems are also
known in the art, and may be used for regulating expression of genes of interest in
the present invention. In some embodiments constitutively active promoters may
be used to express a gene of interest for production of a biosynthetic product.
The biosynthetic product produced by the present invention is typically
an intracellular product, but it may be one or more of the following: i) biomass
from the cultured microorganism; ii) a primary metabolite; iii) a secondary me
tabolite; iv) an extracellular product. The biosynthetic product may be selected
from one or more classes of compounds selected from the list of: alcohols, acids,
enzymes, lipids, antibiotics, proteins, polymers or mixtures thereof. The product
may be the result of genetic engineering of a cell, method of making such are well
known in the art. In genetically engineered cells, production of the product is in
duced by presence of a certain factor in the media, or by temperature shift, or the
product may be constitutively expressed without induction. Further, the product
may be one, the production of which is triggered by certain conditions, such as
presence or lack of specific nutrients, temperature, pH or other. As described
above, the biosynthetic product may be one or more organic compounds produ
ced by a microorganism. In a specific embodiment, the biosynthetic product
comprises any one of lipids, alcohols, enzymes, antibiotics, proteins, polymers,
acids, diacids or any mixture of these.
Specific process conditions such as temperature, nutrient levels, pH
may be provided for microorganisms to produce the wide variety of different or
ganic compounds. These compounds may be considered intracellular products or
extracellular products. Examples of intracellular products may be microbial en
zymes such as: catalase and many others. Recombinant proteins, such as hepatitis
B vaccine, interferon, granulocyte colony-stimulating factor, streptokinase and
others are also intracellular products. However, most enzymes are produced as
extracellular products.
A specific intracellular product by primary metabolism may for exam
ple be single cell protein, which may be produced by many different microorganisms,
such as algae, yeasts, fungi or bacteria. Example of a potential use for this
may include protein additive to feeds for both human and animal use. The production
of intracellular products which are secondary metabolites are also pr e
ferred, and may in non-limiting example encompass microbial oil (e.g. produced
under N, P, S, Fe etc. starvation conditions), wax esters (e.g. produced under N
starvation conditions) and polyhydroxyalkanoate (PHA) (e.g. produced under N,
P, S, Fe etc. starvation conditions), which may be by either aerobic or anaerobic
conditions. Different microorganisms are well known for their ability to produce
either short chain length PHA, or medium chain length PHA). Ergotamine, produ
ced by Claviceps purpurea organisms are also secondary metabolites of interest as
products of the disclosed processes.
Extracellular products include primary metabolites and secondary me
tabolites. Primary metabolites may be produced by aerobic or anaerobic produc
tion. In some instances, a change from anaerobic to aerobic or from aerobic to anaerobic/
microaerobic conditions are needed during production processes.
Extracellular biosynthetic products include but are not limited to for
example insulin, amylase, protease, pectinase, glucose isomerase, cellulase, hemicellulase,
lipase, lactase, streptokinase, microbial oil (e.g. produced under N, P, S,
Fe etc. starvation conditions), isoprenoids (e.g. produced under N starvation con
ditions), ethanol, acetic acid, butyric acid, succinic acid , lactic acid, propionic acid,
malic acid, fumaric acid, pyruvic acid, n-Butanol, iso-butanol, 1,3-propanediol,
1,4-butanediol, ABE or IBE-fermentation, which is a two-stage process and which
is preferred, citric acid (e.g. under Fe-starvation conditions), amino acids, glutam
ic acid (e.g. under nutrient deprivation and low biotin conditions), vinegar (acetic
acid produced from ethanol), 3-hydroxypropionic acid (HPA), 3-hydroxybutyric
acid, 2,3-butanediol, isopropanol, L-isopropanol, itaconic acid (e.g. under Pstarvation
conditions), gluconic acid, adipic acid, microbial lipid, litaconic acid,
enzymes, antibiotics, fatty acid derivatives, terpenoids, hydrocarbons, hydroxyalkanoic
acids (and derivatives thereof), isobutene, and vitamins such as riboflavin.
The carbon sources may be sugars during the growth phase and oils
during the product formation phases. The carbon source may be selected from the
group consisting of sugars (e.g. molasses or lignocellulose derived sugars), fats or
oils or lipids (e.g. waxes, wax esters, sterols, terpenoids, isoprenoids, carotenoids,
polyhydroxyalkanoates, nucleic acids, fatty acids, fatty acid derivates, fatty alco
hols, fatty aldehydes, fatty acid esters, fatty amines, medium and long chain dicarboxylic
acids, epoxy fatty acids, long chain diols and polyols, phospholipids, glycolipids,
sphingolipids or acylglycerols, such as triacylglycerols, diacylglycerols, or
monoacylglycerols), gases (e.g. synthesis gas) and any combinations thereof. However,
the carbon sources suitable for the present invention are not limited to
above group. In different method steps of the present invention the carbon source
may be selected independently from other carbon sources used in other method
steps.
In one embodiment a carbon source or carbon sources are gaseous.
The advantage of the present innovation is that by having high cell densities in
the product formation phase the uptake rate of carbon source from gas phase is
improved and thus the residual carbon source concentration in the fermentation
off gas is lower. This also improves the overall yield of product versus fed carbon
source.
Formation of biosynthetic products by secondary metabolism may be
accomplished by the methods of the present invention. Such products are typical
ly formed during the late exponential or stationary growth phase, and may for
example be selected from the list of enzymes (induction needed), antibiotics, fatty
acid derivatives, isoprenoids, terpenoids etc. hydrocarbons, hydroxyalkanoic
acids (and derivatives thereof, isobutene, gaseous products), and vitamins such as
riboflavin.
In some embodiments of the present invention the biosynthetic prod
uct is one or more lipids selected from the list consisting of: long chain diacids,
hydroxyl fatty acids, long chain diols, fats, oils, waxes, farnesene type products,
wax esters, sterols, terpenoids, isoprenoids, carotenoids, polyhydroxyalkanoates,
nucleic acids, fatty acids, fatty acid derivates, fatty alcohols, fatty aldehydes, fatty
acid esters, fatty amines, medium and long chain dicarboxylic acids, epoxy fatty
acids, long chain diols and polyols, phospholipids, glycolipids, sphingolipids and
acylglycerols, such as triacylglycerols, diacylglycerols, or monoacylglycerols. Mic
roorganisms capable of producing such compounds are well known in the art. The
present invention provides cost effective bioreactor setup allowing for use of
convenient methods for efficient production of these types of compounds.
The present invention also provides a cascade of serially connected bioreactors
or a bioreactor system comprising at least two serially connected b io
reactors and one or more cell concentration devices positioned between and in
line with or in connection to two serially connected bioreactors in said array of
bioreactors. Such bioreactors will in a specific embodiment for aerobic product i
on of biosynthetic products, have means for stirring and aeration, or may be air
lift bioreactors. In some specific embodiments where fermentation is anaerobic,
stirred bioreactors or bioreactors which do not use air lift for agitation of cells
may be utilized. In some other specific embodiments where fermentation is anae
robic and the carbon source(s) is(are) gaseous, air lift type bioreactors (e.g. bub
ble column) may be utilized.
In some embodiments of the present invention the one or more cell
concentration devices are positioned between and in line with each two serially
connected bioreactors in said array of bioreactors.
In some embodiments of the present invention an array of at least
three serially connected bioreactors is provided. Such a bioreactor setup allows
for exposing the cells to at least three different culturing conditions, and allows
the bioreactors to be run in a continuous manner, or in a fed batch mode. This
means that the cells may for example be subjected to conditions for exponential
growth, late exponential growth and production in individual bioreactors.
In some embodiments of the present invention an array of at least four
serially connected bioreactors is provided. Such a bioreactor setup allows for exposing
the cells to at least four different culturing conditions if the bioreactors are
run in a continuous manner. This means that the cells may be subjected to condi
tions for exponential growth, late exponential growth production and production.
In some embodiments, the invention presents a process for production
of biosynthetic products wherein the process comprises the steps presented in
figure 1 (Embodiment A) and figure 2 (Embodiment B :
Embodiment A: A process for production of microbial products or b io
synthetic products, comprising a cascade of steps serially connected to each ot
her, wherein the first step comprises k number of parallel biomass production
reactors, followed by m number of parallel first product formation reactors, followed
by n number of parallel second product formation reactors, and wherein
the first product formation reactors are connected to or in flow connection with
one or more concentration devices positioned to extract microorganism rich fr ac
tion and microorganism poor fraction from the first production reactors and
recycle the microorganism rich fraction back to the first product formation reactors.
Embodiment A.l: a process according to embodiment A, wherein K =
1-x, such as anyone from 1 to 50, such as anyone from 1 to 25, such as anyone
from 1 to 10, such as any one of 1, 2, 3, 4, 5, 6, 7, 8, 9 or 10.
Embodiment A.2: a process according to embodiment A or A.l, wherein
m = 1-y, such as anyone from 1 to 50, such as anyone from 1 to 25, such as
anyone from 1 to 10, such as any one of 1, 2, 3, 4, 5, 6, 7, 8, 9 or 10.
Embodiment A.3: a process according to embodiment A, A.l or A.2,
wherein n = 0-z, such as anyone from 0 to 50, such as anyone from 1 to 25, such as
anyone from 1 to 10, such as any one of 0, 1, 2, 3, 4, 5, 6, 7, 8, 9 or 10.
Embodiment A.4: a process according to any one of embodiments A,
A.l, A.2, or A.3, wherein the individual feeds to the bioreactors k, m and n are of
the same composition.
Embodiment A.5: a process according to any one of embodiments A,
A.l , A.2, or A.3, wherein the individual feeds to the bioreactors k, m and n are of
different composition.
Embodiment B: A process for production of microbial products or biosynthetic
products, comprising a cascade of steps, wherein the first step comp
rises k number of parallel biomass production reactors, followed by m number of
first product formation reactors, followed by n number of second product forma
tion reactors, and wherein one or more concentration devices are positioned
between the biomass production reactors and the first product formation r eac
tors to extract microorganism rich fraction and microorganism poor fraction from
the biomass production reactors and pass the microorganism rich fraction for
ward into the first product formation reactors.
Embodiment B.l: a process according to embodiment B, wherein K =
1-x, such as anyone from 1 to 50, such as anyone from 1 to 25, such as anyone
from 1 to 10, such as any one of 1, 2, 3, 4, 5, 6, 7, 8, 9 or 10.
Embodiment B.2: a process according to embodiment B or B.l, whe
rein m = 1-y, such as anyone from 1 to 50, such as anyone from 1 to 25, such as
anyone from 1 to 10, such as any one of 1, 2, 3, 4, 5, 6, 7, 8, 9 or 10.
Embodiment B.3: a process according to embodiment B, B.l or B.2,
wherein n = 0-z, such as anyone from 0 to 50, such as anyone from 0 to 25, such as
anyone from 0 to 10, such as any one of 0, 1, 2, 3, 4, 5, 6, 7, 8, 9 or 10.
Embodiment B.4: a process according to any one of embodiments B,
B.l , B.2, or B.3, wherein the individual feeds to the bioreactors k, m and n are of
the same composition.
Embodiment B.5: a process according to any one of embodiments B,
B.l, B.2, or B.3, wherein the individual feeds to the bioreactors k, m and n are of
different composition.
In some embodiments according to any one of embodiments A-B5, a
first growth step is optimized for exponential growth, and the second growth step
is for late stage exponential growth.
In some embodiments according to any one of embodiments A-B5, a
first product formation step is for production of biosynthetic product, and the
second product formation step is for production of biosynthetic product under
nutrient starvation.
In some embodiments according to any one of embodiments A-B5, m i
crobial cell status regarding growth is controlled by one or more of nutrients, pH,
temperature, cell density, amount of inhibitory metabolic waste products, carbon
source accessibility, oxygen content in media (in case of aerobic reactor), agita
tion, pressure, flow rate, dilution rate, stirring, inlet, outlet and redox potential.
In some embodiments according to any one of embodiments A-B5, m i
crobial cell status regarding production of biosynthetic products is controlled by
one or more of nutrients, pH, temperature, cell density, amount of inhibitory met
abolic waste products, carbon source accessibility, oxygen content in media (in
case of aerobic reactor), production inducing factors, agitation, pressure, flow
rate, dilution rate, stirring, inlet, outlet and redox potential.
In some embodiments, the present invention is used for microbial oil
production. This has been exemplified in the following:
The cultivation is performed in any kind of bioreactor that is supplied
with air, such as stirred tank bioreactors or air lift bioreactors. Bioreactor types
can be optimized based on the process requirements and different bioreactor
types can be used in one cascade. Typically the biomass production phase (first
bioreactor in a cascade or the biomass production reactor) requires the highest
oxygen transfer rate while oil production phase does not require as high oxygen
transfer rate. According to one embodiment of the invention air-lift bioreactors
are used in the cascade. According to another embodiment of the invention both
stirred tank bioreactors and air-lift bioreactors may be used in the cascade.
In a specific embodiment the bioreactor system comprises at least two
bioreactors which are used in a cascade and a cell concentration step arranged
between the biomass growth and oil production phases for microbial oil production,
(see figure 3). The concentration step may alternatively be positioned after
the production reactor, in which case the cell concentrate is fed back into the pro
duction reactor (figure 4). In a specific embodiment, the bioreactor system com
prises at least three bioreactors, even more specifically at least four bioreactors,
which are used in a cascade with the following set-up and which are operated in
continuous mode (figure 5). The operation of different bioreactors in the cascade
in an exemplified cultivation process for microbial oil production comprising four
bioreactors is additionally described below:
Bioreactor 1 (first bioreactor in a cascade) is used to cultivate micro
organism biomass and is fed with cultivation medium containing adequate
amounts of nutrients allowing efficient biomass growth. Thus, bioreactor 1 is o p
erated at the maximum growth phase of the growth curve. The cell concentration
in bioreactor one is at least lOg/1, more specifically at least 20 g/1, even more sp e
cifically at least 40 g/1. The lipid content of microbial biomass in bioreactor 1 is
typically less than 20% of cell dry weight. Depending on which microorganism is
used, different nutrient content may be present in the media, e.g. typically the
carbon to nitrogen ratio will be less than 100.
Cell concentration, such as centrifugation decanting or filtration, is
performed after the first bioreactor resulting in an increase in biomass concentra
tion as compared to the biomass concentration in the bioreactor from which the
mixture originated (weight % as measured in g/1, cfu (colony forming units), or
by optical density) of at least 50%, more specifically of at least 100%, such as any
one of 10%, 20%, 30%, 40%, 50 %, 60%, 70%, 80%, 90%, or more specifically by
at least 100%.
The purpose of bioreactor 2 (second bioreactor) is to induce nutrient
starvation and lipid production of microorganisms. Thus, the bioreactor is oper
ated at the late exponential phase in the growth curve. Specifically, bioreactor 2 is
operated so that nitrogen (or other nutrient such as. P, Fe or S) is consumed by
microorganisms. The biomass concentration in bioreactor 2 is specifically at least
30 g/L, more specifically at least 60 g/1, more specifically at least 80 g/1. The lipid
content of microbial biomass in bioreactor 2 is typically less than 20% of cell dry
weight.
The purpose of bioreactor 3 (third bioreactor) is to produce oil which
is produced in the secondary metabolism. The bioreactor is operated in the st a
tionary growth phase of the growth curve. Specifically, bioreactor 3 is operated in
the nitrogen (or other nutrient such as. P, Fe or S) starvation to achieve efficient
lipid production. The biomass concentration in bioreactor 3 is specifically at least
50 g/1, more specifically at least 80 g/1, more specifically at least 120 g/1. The lipid
content of microbial biomass in bioreactor 3 is typically more than 30% of cell
dry weight, more specifically at least 40% of cell dry weight.
The purpose of bioreactor 4 (fourth bioreactor) is to produce oil and
to consume residual sugars from the growth medium. The bioreactor is operated
in the (late) stationary growth phase of the growth curve. Specifically, bioreactor
4 is operated in the nitrogen (or other nutrient such as. P, Fe or S) starvation and
addition of carbon source is minimal, i.e. bioreactor 4 is run at near carbon limita
tion. Especially when the growth medium includes mixed sugars, the carbon components,
which are more difficult to use by the microorganisms, are used in the
last bioreactor.
Bioreactor 4 is used to minimize sugar loss by allowing microorgan
isms to utilize residual sugars including sugars that are not preferred by microor
ganisms. The biomass concentration in bioreactor 3 is specifically at least 60 g/1,
more specifically at least 100 g/1, more specifically at least 150 g/1. The lipid con
tent of microbial biomass in bioreactor 4 is typically more than 30% of cell dry
weight, more specifically at least 40% of cell dry weight, even more specifically at
least 50%, even more specifically at least 60%.
After bioreactor 4, microorganism cells are collected from microorganism
culture and the biosynthetic product (e.g. microbial lipid) is recovered. Where
the biosynthetic product is microbial lipid, the overall lipid productivity calculat
ed over the cultivation (over the whole process) according to the invention is at
least 0.3 g/(l*h), specifically at least 0.5 g/(l*h), more specifically at least 1.0 g/(l*h),
more specifically at least 1.5 g/(l*h), more specifically at least 2.0 g/(l*h), even
more specifically at least 2.5 g/(l*h).
When describing the embodiments of the present invention, the com
binations and permutations of all possible embodiments have not been explicitly
described. Nevertheless, the mere fact that certain measures are recited in mutu
ally different dependent claims or described in different embodiments does not
indicate that a combination of these measures cannot be used to advantage. The
present invention envisages all possible combinations and permutations of the
described embodiments.
The terms "comprising", "comprise" and comprises herein are intend
ed by the inventors to be optionally substitutable with the terms "consisting of",
"consist of" and "consists of", respectively, in every instance.
Examples
Example 1 - Rhodosporidiumfluviale. lipid production in fed-batch cu l
tivation
Rhodosporidium fluviale strain CBS 9465 was cultivated in 10 1 scale
bioreactor in glucose fed-batch cultivation for microbial oil production.
For inoculum production following preculture medium was prepared
consisting of (per liter of water): glucose 40 g; yeast extract 5 g; (NH4)2S04 2.5 g;
K2HPO4 0.5 g; KH2PO4 1.0 g; MgS04x7H20 1.0 g; CaCb 2 H2O 0.1 g. Medium was
divided in 50 ml batches and was sterilized by autoclaving at 121°C for 15 min.
Preculture medium was inoculated with 1% inoculum and incubated at 30°C
temperature for 18 hours.
Culture medium composed of per liter of water: glucose 40 g; yeast ex
tract 8 g; (NH4)2S04 3.5 g; KH2PO4 7 g; MgS04 7 H2O 2.5 g; CaCb 2 H2O 0.1 g;
ZnS04 7 H2O 0.0008 g; CuCb 2 H2O 0.00008 g; MnS04 H2O 0.0008 g; FeS04
0.0004 g and NaHP04 0.5 g. Medium was in situ sterilized at 121°C for 15 min.
Culture medium was inoculated with 10% inoculum and after inocula
tion the starting volume was 8 1. The cultivation was done at 30°C temperature.
During cultivation pH was maintained at 6 by adjusting with 3 MNaOH. Aeration
rate was 1 m and pCb (dissolved oxygen) was maintained above 20% using agitation
cascade with max 1000 rpm agitation rate. Struktol J647 was used as antifoam
agent and it was pulsed to the cultivation periodically throughout the cult i
vation. After 23 h cultivation additional 4 g/1 yeast extract was added to the cultu
re medium. During cultivation 54% glucose solution fed to the culture to maintain
glucose concentration at 5-60 g/1. Total cultivation time was 95 h.
During the 95 h cultivation in total 85 g/1 biomass with 61% lipid con
tent (per dry weight) were formed, the overall volumetric lipid productivity being
0.54 g/(lh) at the end of cultivation. Maximum volumetric lipid productivity of
0.56 g/(lh) was achieved after 76.5 h incubation with 75 g/1 cell dry weight con
tent containing 58% lipids.
Example 2 - Rhodosporidium fluviale, lipid production in two step fedbatch
cultivation with interim concentration
Rhodosporidium fluviale strain CBS 9465 was cultivated in two step
glucose fed-batch cultivation with interim concentration for microbial oil produc
tion.
For inoculum production culture medium was prepared consisting of
(per liter of water): glucose 40 g; yeast extract 5 g; (NH4)2S04 2.5 g; K2HPO4 0.5 g;
KH2PO4 1.0 g; MgS04 · 7 H2O 1.0 g; CaCb 2 H2O 0.1 g. Medium was divided in 50
ml batches and was sterilized by autoclaving at 121°C for 15 min. Preculture me
dium was inoculated with 1% inoculum and incubated at 30°C temperature for 18
hours.
Culture medium for biomass production step consisted of (per liter of
tap water): glucose 40 g; yeast extract 8 g; (NH4)2S0 4 2.5 g; MgS0 4 · 7 H20 1.5 g;
KH2PO4 5.0 g; CaCb · 2 H2O 0.1 g. Medium was in situ sterilized at 121°C for 15
min. After sterilization chloramphenicol suspended in ethanol was added to the
medium give chloramphenicol concentration of 0.1 g/1.
Biomass production was done in two parallel bioreactors to obtain
enough bioreactor volume bioreactors were inoculated with 10% inoculum and
after inoculation the starting volume was 10 1in both bioreactors, thus the overall
volume for biomass production was 20 1. Cultivation parameters for biomass production
and concentration steps were: temperature 30°C, pH was maintained at 6
by adjusting with 3 MNaOH, aeration rate was 1 m and the p02 (dissolved oxy
gen) was maintained above 30% using stirring cascade with max 1000 rpm agita
tion rate. Cultivations were fed with 54% glucose solution to maintain glucose
concentration between 5-60 g/1.
After 18 h cultivation cell concentration step was started. Cells were
concentrated by using a cross flow filter (Pellicon 2 Mini with two Durapore 0.45
mih cassettes) connected to the bioreactors. During concentration the parallel cu l
tures were combined in to one bioreactor. Concentration time was 6 h and the t o
tal culture volume after concentration was 7 1.
After the concentration step following supplementing nutrient solu
tions were added to the bioreactor: 7 g MgS0 4 · 7 H2O in 20 ml water; 21 g KH2PO4 in
20 ml water and 34 g Yeast Nitrogen Base without Ammonium Sulphate and Am i
no Acids in 600 ml water.
Cultivation parameters after concentration step were: temperature
30°C, pH was maintained at 6 by adjusting with 3 M NaOH, aeration rate was 1
w m and the p02 was maintained above 20% using stirring cascade with max
1000 rpm agitation rate. Cultivation was fed with 54% glucose solution or dry
glucose to maintain glucose concentration between 5-60 g/1. This lipid production
phase was continued for 67 h. Culture volume at the end of cultivation was 9.5 1.
After 18 h biomass production step the biomass concentration was 29
g/1. During concentration step biomass was concentrated to 95 g/1 concentration.
After the concentration biomass and lipid production were continued and the fi
nal biomass concentration was 148 g/1, of which 42% was lipids. The overall lipid
productivity was 0.60 g/(l*h). Compared to batch cultivation 7% improvement in
volumetric lipid productivity was achieved.
Example 3 - Rhodosporidium fluviale, lipid production in two step fedbatch
cultivation with interim concentration
Rhodosporidium fluviale strain CBS 9465 was cultivated in two step
glucose fed-batch cultivation with interim concentration for microbial oil production.
For inoculum production culture medium was prepared consisting of
(per liter of water): glucose 40 g; yeast extract 5 g; (NH4)2S0 4 2.5 g; K2HPO4 0.5 g;
KH2PO4 1.0 g; MgS0 4x7H 20 1.0 g; CaCb 2 H2O 0.1 g. Medium divided in 100 ml
batches was sterilized by autoclaving at 121°C for 15 min. Preculture media were
inoculated with 1% inoculum and incubated at 29°C temperature for 18 hours.
Culture medium for biomass production step consisted of (per liter of
tap water): glucose 40 g; yeast extract 8 g; (NH4)2S0 4 2.5 g; MgS0 4 · 7 H2O 1.5 g;
KH2PO4 5.0 g; CaCb · 2 H2O 0.1 g. Medium was in situ sterilized at 121°C for 15
min. After sterilization chloramphenicol suspended in ethanol was added to the
medium give chloramphenicol concentration of 0.1 g/1.
Biomass production was done simultaneously in three parallel bioreactors
with working volumes 4, 10 and 10 Lafter inoculation with 15% inoculum.
The overall volume for biomass production was thus 24 1. Cultivation parameters
for biomass production were: temperature 29°C, pH was maintained at 6 by adjusting
with 3 MNaOH, aeration rate was 1 vvm and the p02 was maintained ab o
ve 30% using stirring cascade with max 1000 rpm agitation rate. Cultivations w e
r e fed with 54% glucose solution to maintain glucose concentration between 5-60
g/1-
After 26 h cultivation cell concentration step was started. Cells were
concentrated by using a cross flow filter (Pellicon 2 Mini with three Durapore
0.45 mih cassettes) connected to the bioreactors. During concentration step cult i
vation temperature was maintained at 20°C. Concentration time was 6 h and the
total culture volume combined to one 10 Lbioreactor was 6 Lafter the concentra
tion step.
After the concentration step following supplementing nutrient solu
tions were added to the bioreactor: 7 g MgS0 4 · 7 H2O in 20 ml water; 21 g
KH2PO4 in 20 ml water and 34 g Yeast Nitrogen Base without Ammonium Su l
phate and Amino Acids in 600 ml water.
Cultivation parameters after concentration step were: temperature
29°C, pH was maintained at 6 by adjusting with 3 M NaOH, aeration rate was 1
m and the p02 was maintained above 20% using stirring cascade with max
1000 rpm agitation rate. Cultivation was fed with 54% glucose solution or dry
glucose to maintain glucose concentration between 5-60 g/1. This lipid production
phase was continued for 67.5 h. Culture volume at the end of cultivation was 8.8 1.
After 18 h biomass production step the biomass concentration was 29
g/1. During concentration step biomass was concentrated to 137 g/1 concentra
tion. After the concentration biomass and lipid production were continued and
the final biomass concentration was 189 g/1, of which 69% was lipids.
The maximum productivity was 0.90 g/(l*h) after 78 h cultivation. By
this method 38% increase in volumetric lipid productivity was achieved compared
to batch cultivation.
Example 4 - Rhodosporidium fluviale, lipid production in two step con
tinuous cultivation
Rhodosporidium fluviale strain CBS 9465 was cultivated in two step
continuous cultivation for microbial oil production.
Generalized view of the equipment used in the cultivation is shown in
figure 5. Culture media were prepared as batches in medium sterilization tank
and transferred to media tanks for storage. Microbial cell culturing was started
with a batch phase in biomass production bioreactor to allow cells to propagate
before the start of continuous operation. Similar biomass propagation was not
performed in oil production bioreactor.
After suitable propagation time continuous culturing for biosynthetic
product formation was started. Biomass production bioreactor was continuously
fed with media from media storage tanks. Biomass production bioreactor was
operated in overflow mode and culture was continuous transferred forward via
dip tube.
As the bioreactors used for the demonstration were very different in
size, in order to achieve intended dilution rates in oil production bioreactor, only
portion of the culture from biomass production was directed to oil production.
Rest of the culture from biomass production bioreactor was directed to biomass
waste collection tank. In addition to culture continuously transferred from b io
mass production bioreactor 50% glucose solution was also added continuously to
oil production bioreactor. Similarly to biomass production, oil production was
operated in overflow mode and culture was continuously transferred to collection
tank via dip tube.
For inoculum production culture medium was prepared consisting of
(per liter of water): 40 g glucose, 5 g yeast extract, 2.5 g (NH4 2S0 4, 0.83 g
MgCl2x6H20 , 0.53 g K2HPO4, 1 g KH2PO4 and 0.2 g CaCl2x2H20 . Media components
were dissolved in tap water. Medium divided in 50 ml batches was sterilized by
autoclaving at 121°C for 20 min. Preculture media were inoculated with 1%
inoculum of Rhodosporidium fluviale CBS 9465 yeast and incubated at 30°C temperature
for 27 hours in orbital shaker using 250 rpm agitation rate.
Culture medium for batch biomass propagation consisted of (per liter
of water): 5 g yeast extract, 2 g (NH4)2S0 4, 2 g MgCl2x6H20 , 7 g KH2PO4, 0.4 g
CaCl2x2H20 , 0.03 g MnS0 4xH20 , 0.0003 g ZnS0 4x7H20 , 0.0002 g CuClx2H20 ,
0.0002 g Na2Mo04»2H20 and 0.2 g Struktol J647. Medium volume was 60 1. Components
were dissolved in tap water and the medium was in situ sterilized at
121°C for 30 min. After sterilization 6 kg of sterile 50% glucose syrup was added
aseptically to the medium. Oil production bioreactors were filled with water and
sterilized.
Biomass propagation medium was inoculated with 1.7% preculture.
Cultivation parameters for biomass propagation were: temperature 30°C, pH was
maintained at 5.5 by adjusting with 3 M NH4OH, aeration rate was 1.2 m and
agitation rate was 200 rpm.
After 2 1 h cell propagation in biomass production bioreactor, cont inu
ous two step culturing was initiated by starting the medium feed from medium
storage tanks to biomass production bioreactor, culture transfers from biomass
production to oil production bioreactor and to collection tank, biomass transfer
from oil production bioreactors to collection tank and glucose feed to oil produc
tion bioreactor.
Culture media fed to biomass production bioreactor during continuous
operation consisted of (per liter of water): 54.5 g glucose, 8 g yeast extract, 7 g
KH2PO4, 1.5 (or 1.8) g MgCl2x6H 20 (or MgS0 4x7H 20), 0.5 g (NH4)2S0 4, 0.03 g
MnS0 4xH20 (or MnCl2x4H20), 0.0005 g ZnS0 4x7H20 , 0.0002 g CuClx2H20 , 0.0002
g Na2Mo04»2H20, 0.002 g FeS0 4»7H20 , 0.1 g chloramphenicol, 0.0125.10-3 g biotin,
0.00125 g thiamin HC1, 0.00025 g vitamin B12, 0.00125 g pantothenic acid
and 0.06 - 0.18 g Struktol J647. With the exception of ammonium and magnesium
salts, all minerals, glucose, yeast extract and struktol were dissolved in tap water
and the medium base was in situ sterilized in the medium sterilization tank at
121°C for 30 min forming the medium base. Separate stock solutions were pr epa
red of ammonium and magnesium salts. Solutions were sterilized by autoclaving
at 121°C for 20 min and added aseptically to the sterilized medium base. Vitamins
were dissolved in MQ water, sterile filtered and added aseptically to the sterilized
medium base. Chloramphenicol was dissolved in ethanol and added aseptically to
the sterilized medium base.
Cultivation parameters for biomass production during continuous o p
eration were: flow rate to biomass production bioreactor 11 1/h (dilution rate
0.21 h , temperature 29-30°C, pH was adjusted to 5.0 with 3 M NH40H and
foaming was controlled by automatic and periodic manual antifoam agent addi
tions (struktol J647). During cultivation p0 saturation was maintained above 1%
by manually adjusting agitation, aeration and pressure. Agitation rates were b et
ween 200-460 rpm, aeration 70-80 1/min and overpressure 0-250 mbar. Cultivation
active volume was 52 kg.
Cultivation parameters for oil production during continuous operation
were: dilution rate 0.04 h (only including feed from biomass production step,
not including glucose addition, 100 ml overflow of biomass production was fed to
oil production bioreactor every 15 min), volume approximately 9 1, temperature
29-30°C, pH was adjusted to 4.0 with 3 M NaOH and foaming was controlled by
automatic and periodic manual antifoam agent additions (struktol J647). During
cultivation p0 2 saturation was maintained at 5-20% by manual adjusting of agita
tion and aeration. Agitation rates were between 450-590 rpm, aeration 8 1/min.
Oil production cultivation was fed with sterile 50% glucose solution to maintain
glucose concentration between 5-60 g/1.
Oil productivity was calculated based on the dilution rates.
In biomass production, bioreactor biomass concentration varied b e
tween 18-21 g/1 during continuous cultivation. In oil production bioreactor, b io
mass concentration was 48-58 g/1 and biomass oil content in varied between 31-
40% of CDW. The average overall microbial oil productivity for 145 h period of
continuous operation was 0.98 g/(l*h).
Example 5 - Rhodosporidium fluviale, lipid production in two step con
tinuous cultivation with interim concentration
Culture media were prepared as batches in medium sterilization tank
and transferred to media tanks for storage. Microbial culturing was started with a
batch phase in biomass production bioreactor to allow cells to propagate before
the start of continuous operation. Similar biomass propagation was not perfor
med in oil production bioreactor.
After suitable propagation time continuous culturing was started.
Biomass production bioreactor was continuously fed with media from media storage
tanks. Biomass production bioreactor was operated in overflow mode and
culture was continuous transferred forward via dip tube.
As the bioreactors used for the demonstration were very different in
size, in order to achieve intended dilution rates in oil production bioreactor, only
portion of the culture from biomass production was directed to oil production.
Rest of the culture from biomass production bioreactor was directed to biomass
waste collection tank. During continuous operation oil production bioreactor was
also continuously fed with glucose solution. The feeding rate depended on the su
gar consumption. Similarly to biomass production, oil production was operated in
overflow mode and culture was continuously transferred to collection tank via
dip tube.
After starting of the continuous operation, biomass interim concent ra
tion was started. Membrane filtration unit was connected to oil production b io
reactor. Culture from oil production bioreactor was pumped to membrane filter,
the retentate was circulated back to bioreactor and permeate was removed to co l
lection tank while maintaining continuous operation of the bioreactor.
Culture medium for laboratory inoculum production consisted of (per
liter of water): 40 g glucose, 5 g yeast extract, 2.5 g (NH4)2S0 4, 0.83 g MgCl2x6H20 ,
0.53 g K2HPO4, 1 g KH2PO4 and 0.2 g CaChx2H20. Media components were dissolved
in tap water. Medium divided in 50 ml batches was sterilized by autoclaving
at 121°C for 20 min. Preculture media were inoculated with 1% inoculum of
Rhodosporidium fluviale CBS 9465 yeast and incubated at 29°C temperature for
26 hours in orbital shaker using 250 rpm agitation rate.
Culture media for cell propagation cultivation and continuous culturing
for product formation start up composed of (per liter of water): 60 g glucose,
5 g yeast extract, 4.6 g KH2PO4, 0.03 MnS0 4xH20 , 0.08 CaCl2x2H20 , 0.0008
ZnS0 4x7H20 , 0.00025 CuClx2H20 , 0.0002 g Na2Mo0 4x2H20, 0.0024 g FeS0 4 x
7H20 , 0.19 Struktol, 0.5 g (NH4)2S0 4, 0.5 g MgS04 x 7H20 , 0.8 g MgCk x 6H2O, 0.1 g
chloramphenicol, 0.0000125 g biotin, 0.00125 thiamin HC1, 0.00025 g vitamin
B12, 0.00125 g pantothenic acid. With the exception of ammonium and magne
sium salts, all minerals, glucose, yeast extract and struktol were dissolved in tap
water and the medium base was in situ sterilized at 121°C for 30 min forming the
medium base. Separate stock solutions were prepared of ammonium and magne
sium salts. Solutions were sterilized by autoclaving at 121°C for 20 min and added
aseptically to the sterilized medium base. Vitamins were dissolved in MQ water,
sterile filtered and added aseptically to the sterilized medium base. Chloramphenicol
was dissolved in ethanol and added aseptically to the sterilized me
dium base. Oil production bioreactor was filled with water and sterilized.
Biomass propagation medium was inoculated with 1.4% preculture.
Working volume of the cultivation was 70 1. Cultivation parameters for biomass
propagation culturing were: temperature 29°C, pH was maintained at 5.5 by ad
justing with 3 MNH4OH, aeration rate was 1 m and agitation rate was 250 rpm.
After 2 1 h cell propagation in biomass production bioreactor, cont inu
ous two step fermentation was initiated by starting the medium feed from medi
um storage tanks to biomass production bioreactor, culture transfers from biomass
production to oil production bioreactor and to collection tank, biomass
transfer from oil production bioreactor to collection tank and glucose feed to oil
production bioreactor.
After start-up of continuous cultivation culture media composed of
(per liter of water): 70 g glucose, 8 g yeast extract, 7 g KH2PO4, 0.03 MnS04xH 0 ,
0.08 CaCl2x2H20 , 0.0008 ZnS0 4x7H20 , 0.00025 CuClx2H20 , 0.0002 g
Na2Mo0 4x2H20 , 0.0024 g FeS0 4 x 7H20 , 0.19-0.38 Struktol, 0.5 g (NH4 2S0 4, 0.06-
1.47 g MgS04 x 7H20 , 0-1.16 g MgCl2 x 6H20 , 0.1 g chloramphenicol, 0.0000125 g
biotin, 0.00125 thiamin HC1, 0.00025 g vitamin B12, 0.00125 g pantothenic acid.
Medium preparation was similar to preparation of preculture medium.
Cultivation parameters for biomass production during continuous o p
eration were following. During continuous operation flow rate to biomass pro
duction bioreactor was 14.3 1/h. Temperature was 29°C, pH was adjusted to 5.0
with 3 M NH4OH and foaming was controlled by automatic and periodic manual
antifoam agent additions (struktol J647). During cultivation p0 saturation was
maintained above 20% by manually adjusting agitation, aeration and pressure.
Agitation rates were between 200-400 rpm, aeration 65-70 1/min and overpr es
sure 22-178 mbar. Cultivation active volume varied between 34-66 kg.
Cultivation parameters for oil production during continuous operation
were: dilution rate 0-0.13 h_ , active culturing volume was approximately 7.6 1,
temperature 29°C and foaming was controlled by automatic and periodic manual
antifoam agent additions (struktol J647). During cultivation p0 saturation was
maintained at 5-30% by manual adjusting of agitation and aeration. Agitation r a
tes were between 150-830 rpm, aeration 6-8 1/min. Oil production cultivation
was continuously fed with sterile 40% glucose solution to maintain glucose concentration
between 10-60 g/1.
Concentration step in oil production bioreactor was initiated after 165
h cultivation. Membrane filtration unit (Pellicon 2 Mini with three Millipore Durapore
0.45 mih cassettes) was connected to oil production bioreactor. Culture
from oil production bioreactor was pumped to membrane filter, the retentate was
circulated back to bioreactor and permeate was removed to collection tank while
maintaining continuous operation of the bioreactor. Feed flow rate was adjusted
based on concentration performance. Dilution rate varied between 0-0.13 h .
Periodically the filtration unit was disconnected from the bioreactor system for
cleaning and maintenance.
Oil productivity was calculated based on dilution rates.
In biomass production bioreactor biomass concentration varied b e
tween 19-36 g/1 during continuous cultivation. Similarly in oil production b io
reactor biomass concentration varied between 42-85 g/1 and biomass oil content
varied between 32-45% of CDW. As the filtration system was periodically separated
from the system for maintenance, genuine steady state operation was not a c
hieved. However, on average, when the system was operated with interim con
centration, for 316 h period, the lipid productivity was 1.15 g/(l*h). The average
time included the downtime e.g. caused by filter unit maintenance. If the downti
me is disregarded, the average overall lipid productivity was 1.22 g/(l*h). This
was 20% increase to productivity compared to two step continuous culturing
without interim concentration.

Claims
1. A method for producing a biosynthetic product in a cascade of bioreactors,
the cascade comprising a bioreactor system for biomass production
comprising at least one biomass production reactor, and a bioreactor system for
product formation comprising at least one product formation reactor in flow con
nection with a concentration device, which method further comprises:
a) culturing a microorganism in a biomass production bioreactor by
feeding the bioreactor with nutrient rich culture medium allowing efficient
growth of biomass;
b taking at least part of the microorganism culture from the biomass
production reactor of step a) and feeding it to a product formation reactor contai
ning nutrient depleted medium optimized for formation of the biosynthetic pro
duct;
c producing the biosynthetic product in the presence of the nutrient
depleted product formation medium in the product formation reactor,
wherein the cell mass concentration of the microorganism culture of
the product formation reactor is increased by using a concentration device in flow
connection with the product formation reactor.
2. The method according to claim 1, wherein the concentration device
is arranged between the biomass production reactor and the product formation
reactor.
3. The method according to claim 2, wherein at least part of the micro
organism culture taken from the biomass production reactor is fed into the con
centration device arranged between the biomass production reactor and the
product formation reactor to obtain a microorganism rich fraction, which is fed
into the product formation reactor.
4. The method according to claim 1, wherein the concentration device
is arranged downstream of the product formation reactor.
5. The method according to claim 3 or 4, wherein the method further
comprises taking at least part of the microorganism culture from the product
formation reactor and feeding it into the concentration device arranged between
the biomass production reactor and the product formation reactor or down
stream of the product formation reactor to obtain a microorganism rich fraction,
which is fed into the same or subsequent product formation reactor downstream
in the cascade.
6. The method according to any one of claims 1 - 5, wherein the biomass
production reactor is the last biomass production reactor of the cascade and
the product formation reactor is the first product formation reactor of the casca
de.
7. The method according to any one of claims 1 - 6, wherein the bior e
actor system for biomass production comprises at least 2, 3, 4, 5, 6, 7, or 8 bioreactors.
8. The method according to any one of claims 1 - 6, wherein the bior e
actor system for product formation comprises at least 2, 3, 4, 5, 6, 7, or 8 bioreactors.
9. The method according to any one of claims 1 - 8, wherein a microor
ganism rich fraction is obtained by using means or methods selected from the list
consisting of: centrifugation, filtration, settling, flocculation, flotation, decanting.
10. The method according to any one of claims 1 - 9, wherein the bioreactor
system or the cascade of serially connected bioreactors is operated in a
continuous or fed-batch manner or in a combination thereof.
11. The method according to any one of claims 1 - 10, wherein the
method further comprises:
one or more biomass production bioreactors of the cascade are optimised
for microorganism growth at late stage exponential phase, and/or
the product formation bioreactor subsequent to the biomass producti
on bioreactor is optimised for biosynthetic product formation at a stationary
growth phase, optionally under nutrient starvation conditions, and/or
any additional product formation bioreactor is optimised for biosynthetic
product formation with minimal carbon source addition.
12. The method according to any one of claims 1 - 11, wherein the m i
croorganism rich fraction is obtained at a temperature between 3 to 30°C lower
than the bioreactor from which the microorganism culture was taken from
and/or wherein the temperature of the microorganism culture to be fed into the
concentration device is lowered by 3 to 30°C.
13. The method according to any one of claims 1 - 12, wherein the b io
synthetic product is one or more of the following:
i biomass from the cultured microorganism;
ii a primary metabolite;
iii a secondary metabolite;
iv an intracellular product;
v an extracellular product.
14. The method according to any one of claims 1 - 13, wherein the biosynthetic
product is an intracellular product.
15. The method according to any one of claims 1 - 14, wherein the biosynthetic
product is one or more lipids selected from the list consisting of: long
chain diacids, hydroxyl fatty acids, long chain diols, lipids, fats, oils, waxes, farnesene
type products, wax esters, sterols, terpenoids, isoprenoids, carotenoids, polyhydroxyalkanoates,
nucleic acids, fatty acids, fatty acid derivates, fatty alcohols,
fatty aldehydes, fatty acid esters, fatty amines, medium and long chain dicarboxylic
acids, epoxy fatty acids, long chain diols and polyols, phospholipids, glycolipids,
sphingolipids and acylglycerols, such as triacylglycerols, diacylglycerols, or
monoacylglycerols.