Recombinant protein production by engineered Escherichia coli in a pressurized airlift bioreactor: A techno-economic analysis

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Chemical Engineering and Processing xxx (2015) xxx–xxx

Contents lists available at ScienceDirect

Chemical Engineering and Processing: Process Intensification journal homepage: www.elsevier.com/locate/cep

Recombinant protein production by engineered Escherichia coli in a pressurized airlift bioreactor: A techno-economic analysis Gilson Campania , Maurício Possedente dos Santosa , Gabriel Gonçalves da Silvaa , Antônio Carlos Luperni Hortaa,b , Alberto Colli Badinoa,b , Roberto de Campos Giordanoa,b , Viviane Maimoni Gonçalvesc , Teresa Cristina Zangirolamia,b,* a

Chemical Engineering Graduate Program, Federal University of São Carlos (PPEQ-UFSCar), Brazil Department of Chemical Engineering, Federal University of São Carlos, Rodovia Washington Luís, km 235, CEP 13565-905 São Carlos, SP, Brazil c Biotechnology Center, Butantan Institute, Av. Vital Brasil, 1500, CEP 05503-900, São Paulo, SP, Brazil b

A R T I C L E I N F O

A B S T R A C T

Article history: Received 30 June 2015 Received in revised form 24 October 2015 Accepted 27 October 2015 Available online xxx

Standard cultivation protocols for high cell-density cultures of recombinant E. coli rely on air enrichment with oxygen to prevent oxygen limitation. This way of increasing oxygen supply impacts on process economics due to the high cost of pure oxygen. Reactor pressurization is an alternative approach to improve oxygen mass transfer. In the present study, the performance of pressurized airlift bioreactor in growing rE. coli cells was evaluated. Experiments were performed in a pressurized internal-loop airlift bioreactor (ALB) and in a non-pressurized stirred tank reactor (STR), both with 5 L working volume, equipped with a system for automatic control and monitoring of pressure, temperature, pH, and dissolved oxygen. Pressurization showed to be crucial to improve ALB performance in biomass formation (29 gDCW L
1) and protein production (201 mgprotein gDCW
1), resulting in protein productivity (240 mgprotein
1 h
1) and energy efficiency (11 gprotein kWh
1) similar to the achieved in STR (200 mgprotein L
1 h
1 and 13 gprotein kWh
1, respectively). Due to the high cost of pure oxygen, air enrichment revealed to be economically unfeasible for ALB. By pressurizing the bioreactor up to 0.41 MPa, without pure oxygen supply, an 8.7-fold increase in economic efficiency is estimated, what shows the potential of this innovative strategy for aerobic cultures. ã 2015 Elsevier B.V. All rights reserved.

Keywords: E. coli PspA Pressurized airlift Dissolved oxygen tension

1. Introduction In the past four decades rapid advances in the field of molecular biology, microbiology and bioprocess resulted in increased productivity and quality of several biotechnological products [1]. In this context, E. coliplays a special role as one of the most used hosts for the production of heterologous proteins, due to the extensive knowledge gathered about its metabolism and genetics, associated with the availability of accurate and rapid methods for the modification of its genome [2]. Regarding the types of reactors used in E. coli cultures, most of the studies employ conventional bioreactors (stirred tank reactors, STR). In recent years, there has been an increasing interest in the development of non-conventional bioreactors, such as the rockingmotion type [3]. Nevertheless, less attention has been paid to the

* Corresponding author. Fax: +55 16 3351 8266. E-mail address: [email protected] (T. Cristina Zangirolami).

cultivation of E. coli in pneumatic bioreactors such as the airlift type, which are attractive for several biological processes [4–7]. Airlift bioreactors exhibit advantages such as: simplicity of construction, low operational cost, reduced risk of contamination and efficient gas–liquid dispersion with low power consumption [8]. On the other hand, airlift bioreactors present lower oxygen transfer due to the absence of mechanical agitation by impellers. This drawback, however, can be overcome by pressurizing the reactor, thus increasing oxygen solubility in the liquid phase and hence the oxygen transfer rate [9]. Moreover, previous studies have reported that the pressurization of pneumatic reactors also increases gas hold-up and the volumetric oxygen transfer coefficient [10–15]. To date, there are few studies investigating the cultivation of E. coli in pressurized reactors. Anyway, improvements on biomass productivity were observed for conventional bioreactors [16–19]. In addition, increasing reactor pressure enhances the energy and cost efficiencies of the dissolved oxygen control system [20]. It can be thus seen that the manipulation of bioreactor working pressure poses as a promising approach to intensify aerobic bioprocess,

http://dx.doi.org/10.1016/j.cep.2015.10.020 0255-2701/ ã 2015 Elsevier B.V. All rights reserved.

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Nomenclature

ALB C Cx_final CPspA CPspA_final EEX EEPspA Di Dt DOT fc HL N p p0 p1 p2 P0 PC PS PspA PPspA Px Q AIR Q GAS Q O2 R STR T tf in V yO2 YPspA/X

hC hS g

$OV

inernal-loop airlift bioreactor cell concentration (gDCW L
1) final cell concentration (gDCW L
1) PspA concentration (gPspA L
1) final PspA concentration (gPspA L
1) economic efficiency in terms of biomass (gDCW US $
1) economic efficiency in terms of PspA (gPspA US$
1) impeller diameter (m) tank diameter (m) dissolved oxygen tension (percent saturation) correction factor (dimensionless) liquid height (m) stirrer speed (s
1) absolute pressure (Pa) absolute pressure in the simulated experiment (Pa) inlet absolute pressure of the compressor (Pa) outlet absolute pressure of the compressor (Pa) stirrer power consumption for the non-aerated system (W) compressor power consumptions (W) stirrer power consumptions for the aerated system (W) pneumococcal surface protein A productivity of PspA (mg L
1 h
1) productivity of biomass (gDCW L
1 h
1) volumetric air flow rate (m3 s
1) volumetric gas flow rate (m3 s
1) volumetric oxygen flow rate (m3 s
1) universal gas constant (m3 Pa K
1 mol
1) stirred tank reactor temperature (K or  C) final time of cultivation (s) working volume of the reactor (m3) oxygen mole fraction (dimensionless) specific production of PspA (mg gDCW
1) degree of efficiency of the compressor degree of efficiency of the stirrer isentropic exponent (dimensionless) overall cost related to the DOT control system (US $)

improving productivity, efficiency and sustainability. However, so far, this alternative technology has not been properly investigated in pneumatic bioreactors, especially in terms of energy and economic efficiencies. Therefore, the aim of this study is to assess the performance of a pressurized internal-loop airlift bioreactor (ALB) when challenged to meet high oxygen demand featured by recombinant E. coli cultures. Usual performance indices, such as biomass and protein productivities, in combination with energy and cost efficiency analyses were used to compare ALB and stirred tank reactor (STR) and to evaluate the proposed process intensification strategy. 2. Materials and methods 2.1. Microorganism and culture media Experiments were carried out with E. coli BL21(DE3) harboring the plasmid pET37b(+)/PspA4Pro. These cells produce a

recombinant fragment of pneumococcal surface protein A (PspA) from clade 4, containing the N-terminal alpha-helical domain and the first block of the proline rich region [21] under control of lacUV5 and T7lac promoters. Due to its remarking immunogenic characteristics, PspA is a promising candidate as carrier protein in conjugate vaccines against Streptococcus pneumoniae [22] or as an alternative serotype-independent pneumococcal vaccine [23]. The clone was kindly provided by Dr. Eliane Miyaji from the Laboratory of Molecular Biology, Butantan Institute (São Paulo, Brazil). The sterile auto-induction culture medium [24] contained glycerol (60 g L
1), glucose (10 g L
1), lactose (20 g L
1), yeast extract (5 g L
1), phytone (10 g L
1), KH2PO4 (3.4 g L
1), NH4Cl (2.7 g L
1), Na2SO4 (0.7 g L
1), ferric citrate (100.8 mg L
1), CoCl26H2O (2.5 mg L
1), MnCl24H2O (15 mg L
1), CuCl22H2O (1.5 mg L
1), H3BO3 (3 mg L
1), Na2MoO42H2O (2.1 mg L
1), Zn (CH3COOH)2H2O (33.8 mg L
1), EDTA (14.1 mg L
1), MgSO47H2O (0.5 g L
1), thiamine (45 mg L
1), kanamycin (100 mg L
1), and polypropylene glycol (0.03% v v
1). This culture medium is used for batch mode, which in turn is suitable for airlift bioreactors since they are not adequate to carry out fed-batch processes [25]. This complex medium promotes fast growth, intense biomass formation and it showed to be suitable for recombinant protein production by E. coli [26,27]. Inoculum medium was prepared with 5.0 g L
1 glycerol, 0.5 g L
1 glucose (without lactose). 2.2. Bioreactors Internal-loop airlift bioreactor (ALB) and a stirred tank reactor (STR), both with 5 L working volume, were used in this study. ALB has a glass window on its stainless steel body and a cross-piece type sparger, as described by Campani et al. [9]. STR consists on a glass and stainless steel (thermal jacket) vessel with three Rushton turbine impellers [28]. The bioreactors were monitored by SuperSys_HCDC1 software [28]. On-line data was logged via Compact FieldPoint (National Instruments, model cFP2020). The experimental setup is shown in Fig. 1. 2.3. Experimental procedure Two cultivations were carried out in the ALB (pressurized and non-pressurized conditions) and one in the STR (non-pressurized). All experiments were started by picking up cells from a glycerol stock suspension at
80  C and transferring them to a LuriaBertani (LB) agar plate supplemented with 100 mg L
1 kanamycin, which was incubated at 37  C for 24 h. A single colony was then inoculated in 0.5 L flask containing 50 mL of inoculum medium (pre-inoculum) that was later kept in a controlled environment incubator shaker (New Brunswick) at 37  C and 270 rpm until optical density (OD) of 2.5 (at 600 nm). Inoculum was prepared by transferring 12 mL of the pre-inoculum to 300 mL of fresh inoculum medium in 1 L flask. The inoculum was then incubated at 37  C and 270 rpm until an optical density of 2.5, when it was completely transferred (300 mL) to the bioreactor containing 4.7 L of fresh medium. During the bioreactor cultivations, pH was maintained (on/off control) at 6.7 (growth phase) and 6.9 (induction phase) by addition of NH4OH (7% w w
1) or HCl (9% w w
1) solutions. Induction phase starts after glucose depletion, when the consumption of lactose (inducer) is initiated by the cells [24]. Glucose depletion was detected from on-line data as a sharp decline on carbon dioxide evolution rate and confirmed using glucose oxidase assay (GOD-PAP, Laborlab, Brazil). Temperature was controlled (in the range of 15–31  C, depending on DOT control strategy used) by manipulating the temperature of the water re-circulating in the bioreactor jacket. Dissolved oxygen tension (DOT) was determined by an amperometric probe (Mettler Toledo, model InPro 6830) and

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3

Fig. 1. Experimental setup: (1) gas mass flow controllers (Aalborg, model GFC), (2) gas valve actuator, (3) temperature, pressure, dissolved oxygen tension, and pH sensors, (4) peristaltic pumps for pH control (Masterflex, model C/L), and (5) Compact FieldPoint.

different DOT controllers were employed according to the bioreactor type. The heuristic algorithm for DOT control in STR, developed by Horta et al. [28], is based on the manipulation of three variables (stirrer speed and air and oxygen flow rates) to maintain DOT at 30% saturation (relative to saturation in the calibration procedure at 0.1 MPa and 31  C). For ALB, DOT was kept between 20 and 50% of saturation using four manipulated variables: air and oxygen flow rates (steps of D, with 0.1n  D (L min
1)  0.5), temperature (steps of 0.1  C) and absolute pressure up to 0.25 MPa (steps of 0.01 MPa). The chosen DOT range was based on previous experiments (data not shown) which demonstrated that there is not drift towards fermentative metabolism as long as DOT is maintained above 10% of saturation. Pressure was controlled employing a pressure sensor (SMC, model ISE30A-01-F) and a gas valve actuator (SMC, model VY1100), both placed in the exit gas stream. The maximum operating pressure was fixed at 0.25 MPa to comply with the specifications of the peristaltic pumps used to control the pH. For the cultivation carried out under overpressure, the volumetric gas flow rate (Q GAS ) was held constant within the _ proportionally bioreactor by changing the molar gas flow rate (n) to the pressure, according to the ideal gas law (Eq. (1)). Q GAS ¼

RT  n_ p

ð1Þ

where T (K) is the temperature, p (Pa) is the absolute pressure, and R (m3 Pa K
1 mol
1) is the universal gas constant. 2.4. Analytical methods Cell concentration (CX) was assessed by culture broth optical density measurements (l = 600 nm) and dry cell weight (gDCW L
1). Metabolites (acetate, lactate and formate) and carbon source concentrations were analyzed by HPLC using an Aminex HPX-87H column (Bio-Rad) and a 5 mM sulfuric acid solution as mobile phase (0.6 mL min
1). The temperature was 60  C. Carbon sources were measured with a refraction index detector (Waters 410), while organic acids were detected at 210 nm (Waters 486 UVdetector).

Cell disruption was performed by sonication using 10 pulses of 30 s (26 W) at 20 kHz with intervals of 30 s. The total soluble protein from cell extracts was quantified using the Bradford assay [29]. The expression of soluble PspA was analyzed by 12% SDSPAGE [30]. Gels were stained with Coomassie brilliant blue R and photographed to estimate protein production by densitometry using the ImageJ1 software [31]. To follow plasmid loss during the cultivations, diluted samples (1:107) of culture broth were spread aseptically onto plates of LB agar and incubated for 24 h at 37  C. From each sample, more than fifty colonies were transferred to new plates with and without 100 mg L
1 kanamycin, which were further incubated for 24 h at 37  C. Plasmid stability of each sample was assessed by calculating the percentage of antibiotic resistant colonies in relation to the number of colonies formed in plates without antibiotic. 2.5. Estimation of economic and energy efficiencies Energy efficiency was determined as the ratio between final biomass or PspA production and the total power consumed by the DOT control system, comprising the energy spent by both compressor and stirrer (STR) or just by the compressor (ALB). The compressor power consumption was calculated assuming gas compression under isentropic condition, in a single stage with degree of efficiency (hC ) of 0.7, and considering air as a perfect gas (Eq. (2)) [20]. Total pressure loss of 0.1 MPa was considered. The stirrer power consumption was estimated by Rushton’s method [32,33], combined with the correlation developed by Michel and Miller [34] for gas–liquid agitated systems (Eq. (3)). The degree of efficiency of the stirrer (hS ) considered was 0.65 [20]. The correction factor (f c ) (Eq. (4)) proposed by Aiba et al. [35] was used to take into account geometry differences with respect to the reactor employed in the Rushton’s method [32–33]. "
g
1 # 1 g p2 g  p1  PC ¼  Q AIR 
1 ð2Þ hC p1 g
1

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ð3Þ

sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi ðDt =Di Þ  ðHL =Di Þ fc ¼ ðDt =Di Þ  ðHL =Di Þ

ð4Þ

where PC and PS (W) are the compressor and stirrer power consumptions, Q AIR and Q GAS (m3 s
1) are the air and gas flow rates, p1 and p2 (Pa) are the inlet and outlet absolute pressure of the compressor, g is the isentropic exponent equal to 1.4 (nitrogen and oxygen gases), P0 (W) is the stirrer power consumption for the non-aerated system estimated by Rushton's method [32–33],N (s
1) is the stirrer speed, Di (0.078 m) and Dt (0.16 m) are the impeller and tank diameters, HL (0.267 m) is the liquid height, and ‘*’ indicates the geometry of the STR used in the current study. The ratios (Dt =Di ) and (HL =Di ) for a standard geometry bioreactor are equal to 3. The overall economic efficiency (EE in g US$
1) of the DOT control system with respect to biomass and PspA production for each experiment was estimated considering the power consumption costs of the compressor and the stirrer (0.093 US$ kWh
1 [36]). The cost of pure oxygen fed to the bioreactors was taken as 0.47 US$ std-m
3 of pure oxygen [37], where std-m3 is m3 stated for mass controller reference conditions (i.e. 21.1  C at 1 bar), as described by Eqs. 5–7. tfin Z

$OV ðUS$Þ ¼ 0:093 

  C
1 ¼ X EEX gDCW US$

V

 CX $OV

final

(a)

V

ð7Þ

50

9 8

-1

7 30

6

20

5 4

-1

lactose (g.L )

10

3 2

0

1

-10

-1

DOT(%)

10

90

(b)

4

8

12

16

20

Time (h)

24

28

32

36

1.0

50

0.8

40

60 DOT

30 0

0.6

T

0.4

0.2 QAIR

30 20

0

3

10 0

0.0

QO2

0 0

inducon phase

-1

11

formic acid (g.L )

inducon phase

60

acec acid (g.L )

-1

biomass (gDCW.L )

final

40

glucose (g.L )

-1

glycerol (g.L )

ð6Þ

$OV

  Y PspA
1 ¼ EEPspA gPspA US$

-1

Two experiments were performed in the airlift bioreactor (ALB), one at atmospheric pressure and another with pressurization up to 0.25 MPa. In the stirred tank reactor (STR), one cultivation was carried out at atmospheric pressure as reference condition. The experiment in the ALB at atmospheric pressure (Experiment 1) showed a maximum biomass concentration of 20 gDCW L
1 and specific production of PspA (YPspA/X) of 47 mg gDCW
1 after 35 h of cultivation, when it was interrupted due to insufficient oxygen transfer rate - Fig. 2(a) -. Formic and acetic acids accumulated before glucose depletion (up to 0.84 and 1.45 g L
1, respectively) were partially consumed after that, simultaneously with glycerol. Lactic acid formation was not detected throughout the cultivation. Plasmid stability was high during the induction phase (superior to 90%), indicating a moderate metabolic stress in the cells (possibly related to the reduced temperature during induction phase, which was lower than 21  C). The dynamics of dissolved oxygen tension (DOT) throughout the experiment is shown at Fig. 2(b), together with the changes in the temperature and in the flow rates of air and oxygen. As shown in Fig. 2(b), DOT was controlled between 20 and 50% of saturation in a cascade mode, firstly manipulating air flow rate (Q AIR ), then oxygen flow rate (Q O2 ) and finally temperature (T) (minimum of 15  C). Maximum gas flow rate (Q AIR þ Q O2 ) employed was 0.33 L s
1 in order to avoid excess foaming, which

0

final

ð8Þ

3. Results and discussion

ð5Þ

Q O2 dt

yO2 p 0:21

where p0 (Pa) is the pressure in the simulated experiment, p (Pa) is the pressure in the pressurized ALB (Experiment 2), and yO2 is the oxygen mole fraction of the enriched inlet gas.

tfin Z

ðPC þ PS Þdt þ 0:47  0

YPspA/X (mgPspA.gDCW)

p0 ¼

T(°C)

Q 0:56 GAS

-1

P20  N  D3i

QO2(L.s )

hS

0:706 

-1

1

QAIR(L.s )

PS ¼

where $OV (US$) is the overall cost related to the DOT control system, C X final (gDCW L
1) is the final cell concentration, Y PspA final ( mg gDCW
1) is final specific production of PspA, V (m3) is the working volume of the reactor, and tfin (s) is the final time of cultivation. The economic efficiencies of the pressurized ALB up to 0.41 MPa (without oxygen enriched gas) were simulated assuming the same partial oxygen pressure (same oxygen transfer rate) achieved in the pressurized ALB up to 0.25 MPa, using Eq. (8)

!0:45

6

9

12

15

18

21

30

33

36

Time (h)

Fig. 2. (a) Substrate consumption, biomass, metabolites and recombinant protein formation. (b) Temperature (T), air and oxygen flow rates (and, respectively), and DOT during cultivation of E. coli in the non-pressurized ALB (Experiment 1). The and values are presented at the reference condition of 0.1 MPa and 21.1  C. Error bars for biomass and YPspA/X are standard deviations from triplicates. Data loss between 24 and 29 h in the graph (b) (x-axis break) is due to a failure in the automatic acquisition system (without affecting the DOT control).

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4

20

-1

160 140 120

3 10

2

1.8

1.4

T

1.2 1.0 0.8 QAIR

100

P

80

QO2

0 0

4

8

12

16

20

36 -1

1.6

32

0.6 0.4

28 24 20

0.2

1

0

40

2.0

DOT

T(°C)

5

YPspA/X (mgPspA.gDCW)

30

acec acid (g.L )

7 6

-1

180

8

inducon phase

QO2(L.s )

200 9

50

(b)

-1

60

100 80 60 40 20 0

DOT(%)

10

40 glucose (g.L ) -1 biomass (gDCW.L )

220

11

QAIR(L.s )

inducon phase

P (MPa)

(a)

-1

glycerol (g.L ) -1 lactose (g.L )

70

5

0

24

3

6

9

12

15

18

21

0.0

16

24

Time (h)

Time (h)

Fig. 3. (a) Substrate consumption, biomass, metabolite and recombinant protein formation. (b) Temperature (T), absolute pressure (p), air and oxygen flow rates (and, respectively), and DOT during cultivation of E. coli in the pressurized ALB (Experiment 2). The and values are presented at the reference condition of 0.1 MPa and 21.1  C. Error bars for biomass and YPspA/X are standard deviations from triplicates.

takes place at higher gas flow rate. The observed DOT disturbances can be attributed to the manipulation of temperature, which was unable to provide a fine DOT control due to its indirect effect on DOT dynamics (reducing the temperature implies a lower cell metabolism and hence lower oxygen uptake rate). With the ALB pressurization (Experiment 2), the bioprocess was considerably improved as shown in Fig. 3(a). A maximum cell concentration of 30 gDCW L
1 was achieved after 23.5 h cultivation and YPspA/X reached 201 mg gDCW
1, a 4.3-fold increase compared to protein production achieved in Experiment 1. Acetic acid concentration accumulated up to 0.61 g L
1and was totally consumed after glucose depletion. Formic and lactic acids were not detected. Plasmid stability dropped from 87% (at beginning of induction phase) to a minimum of 41% after 9 h of induction. This decrease in the percentage of plasmid-carrying cells can be attributed to the metabolic stress triggered by the overexpression of the recombinant protein (PspA), which was higher than that in the non-pressurized ALB. As it can be seen from Fig. 3(b), DOT was controlled automatically in the proposed cascade mode. Firstly, Q AIR was increased up to 0.33 L s
1, when the system started to be gradually pressurized up to the limit of 0.25 MPa. With the system pressurization, mass flow rate of air was proportionally raised, according to the ideal gas law (Eq. 1), in order to keep constant the volumetric gas flow rate within the vessel. When the pressure

3

20

2 10

1 0

0 0

2

4

6

8

10

12

Time (h)

14

16

18

20

40 20 0

DOT

500

4

400

3

300

2

N

200

QAIR

100

1 QO2

0 0

140

60

120 100

2

4

6

8

10

12

14

16

18

80 40

DOT (%)

4

-1

30

600

5

2

60

700

160 -1

-1

80

180

7

800 6

5 Lacc Acid (g.L )

-1

6

100

10 x QO2(L.s )

7

8

2

50

inducon phase

900

N (rpm)

8

(b)

-1

9

1000

YPspA/X (mgPspA.gDCW)

60

Acec Acid (g.L )

10

40

glucose (g.L ) -1 biomass (gDCW.L )

120

11

-1

inducon phase

10 x QAIR(L.s )

(a)

-1

glycerol (g.L ) -1 lactose (g.L )

70

reached its maximum, the inlet gas was enriched with oxygen up to 16% v v
1 (pure oxygen:total gas). Finally, temperature was slowly decreased to a minimum of 26  C in approximately 0.6  C h
1 to slow down oxygen uptake rate. Further decreases in the temperature were not necessary because of the higher oxygen transfer rate achieved with the system pressurization in this experiment. On average, DOT was held above 20% of saturation and the observed disturbances are related to the temperature changing, as reported for Experiment 1. E. coli cultivation was carried out in the non-pressurized STR (Experiment 3) as well. The temperature in this experiment was reduced from 31  C (growth phase) to 25  C (induction phase). Induction phase was performed at 25  C because it is similar to the temperature range employed in the ALB cultivations (Experiments 1 and 2). The main results obtained in this experiment are shown in Fig. 4(a). Final values of Cx and YPspA/X were 39 gDCW L
1 and 105 mg gDCW
1, respectively, after 20 h of cultivation. Low accumulation of metabolites was observed, up to 0.9 g L
1 (acetic lactic) and 0.31 g L
1 (lactic acid), before glucose depletion. After that, these acids were completely consumed. Formic acid was not detected during the cultivation. Plasmid stability was superior to 94% during the induction phase due to the low temperature employed in this phase (25  C), which decreased the expression of the recombinant protein and hence the metabolic burden related

20 0

0 20

Time (h)

Fig. 4. (a) Substrate consumption, biomass, metabolites and recombinant protein formation. (b) Temperature (T), stirrer speed (N), air and oxygen flow rates (and, respectively) and DOT during cultivation of E. coli in the non-pressurized STR (Experiment 3). The and values are presented at the reference condition of 0.1 MPa and 21.1  C. The temperature was 31  C in the growth phase and 25  C in the induction phase. Error bars for biomass and YPspA/X are standard deviations from triplicates and duplicates, respectively.

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Table 1 Performance indices (biomass and PspA productivities) of the E. coli cultivation in the non-pressurized ALB and STR, and in the pressurized ALB (ALB-P).

ALB ALB-P STR

PX (gDCW L
1 h
1)

PPspA (mgPspA L
1 h
1)

0.57 1.27 1.9

27 240 200

-1

Enerngy efficiency (gPspA.kWh )

to it. The stability of the DOT control, as shown in Fig. 4(b), can be attributed to the fine control provided by the stirrer speed manipulation synchronized with the gradual increase in the air flow rate (from 6 to 8 h) and the activation of pure oxygen supplementation (last 2 h) [28]. The main performance indices of Experiments 1, 2 and 3 are displayed in Table 1. System pressurization intensifies the bioprocess in ALB, increasing PspA productivity from 27 mg L
1 h
1 (nonpressurized ALB) to 240 mg L
1 h
1 (pressurized ALB), which is superior to the value obtained in the non-pressurized STR (200 mg L
1 h
1). Concerning biomass productivity, bioreactor pressurization led to a 2-fold increase (from 0.57 to 1.27 gDCW L
1 h
1), getting closer to the value achieved with the STR (1.9 gDCW L
1 h
1). Using the data collected during the Experiments 1, 2 and 3 along with Eqs. (2–4), it is possible to assess the impact of the DOT control strategy on the energy consumption profiles for bioproduct formation. Fig. 5 shows that, for the same range of biomass concentrations (below 25 gDCW L
1), ALB-P control system provided the highest energy efficiency, due to both lower power consumption and increased protein production. The results of energy efficiency combined with the oxygen consumption data can be further expressed as economic efficiency (Eqs. (5–7)). These results are summarized at Table 2. It can be seen that E. coli cultivation in the ALB or ALB-P has lower economic efficiency (related to the DOT control system) with respect to both biomass and PspA production. This inferior performance is mainly due to the higher average flow rate of pure oxygen employed in the ALB (with and without pressurization), which contributes to more than 92% of overall DOT control cost (Table 2). One way to increase the economic efficiency of the ALB, therefore, is to pressurize even more this reactor (more than 0.25 MPa) in order to decrease the amount of pure oxygen gas required to maintain DOT at the desired levels. This strategy was simulated (Eq. (8)) and included in Table 2. The economic efficiency of the pressurized ALB would be significantly increased by pressurizing the bioreactor up to

14

14

12

12

10

10

8

8

6

6

4

4

2

2

ALB ALB-P STR

0 0

5

10

15

20

25

30

35

Table 2 Economic performance indices (overall economic efficiencies EEX and EEPspA) associated to the DOT control system for E. coli cultivation in the non-pressurized ALB and STR, and in the pressurized ALB (ALB-P).

ALB ALB-P STR ALB-Pb

EEX (gDCW US$
1)

EEPspA (gPspA US$
1)

72.96 47.43 1,272.70 414.56

3.43 8.64 133.61 75.51

Cost contributionsa kWh (%)

O2 (%)

6 7.8 95.2 100

94 92.2 4.8 0

a Relative contributions of cost components power and pure oxygen consumptions. b Simulated economic efficiencies assuming the same oxygen transfer rate achieved in the ALB-P, but with system pressurization up to 0.41 MPa instead of enriching the inlet gas with pure oxygen.

0.41 MPa, instead of using pure oxygen gas. This simulated efficiency, however, would still be lower than the observed in the STR. It is important to point out that DOT control in ALB is a research field under development, while oxygen transfer in STR is a rather established issue. The controller structure has a strong influence on process performance and further studies of different configurations can lead to improved productivities and efficiencies. Some promising changes to be tested include modifications in the logics of the cascade controller from ["Qair ! "Pressure ! "QO2 ! #Temperature] to ["Pressure ! "Qair], excluding the temperature gradual decrease (to improve biomass and protein productivities) and O2 enrichment (to reduce the cost) from DOT cascade controller, so that bioreactor overpressure will be left as the main actuator. Furthermore, the economic optimization of DOT control can also be carried out.

4. Conclusions The airlift bioreactor performance when employed for cultures of high oxygen uptake organisms, such as E. coli, was evaluated in several aspects. We showed that bioreactor pressurization is crucial to achieve competitive productivity and energy efficiency values and approach conventional stirred tank bioreactor performance. Furthermore, bioreactor pressurization poses an alternative of lower cost to overcome oxygen limitation during E. coli high density cultures, in comparison to the usual solution, based on air enrichment with pure oxygen. Even for stirred tank bioreactor, the supply of pure oxygen contributes to approximately 4.8% of overall dissolved oxygen controller cost. Its replacement by a bioreactor pressurization system could reduce by half the overall cost related to oxygen transfer. The outcomes of this work have important implications for developing novel strategies for DOT control in pneumatic bioreactors based on gas flow rate and overpressure as manipulated variables. This research will serve as a base for future studies with cultivation of different microorganisms in pressurized airlift bioreactors, contributing to the development of cost-effective bioprocess. Acknowledgements

0 40

-1

Cx (gDCW.L ) Fig. 5. Energy efficiency of DOT control system (expressed as the ratio protein production/power consumption) during E. coli cultivation in the non-pressurized ALB and STR, and in the pressurized ALB (ALB-P).

The authors would like to thank FAPESP (The State of São Paulo Research Foundation) (Grant Processes 2008/05207-4 and 2011/ 16605-3) and CAPES (Brazilian Federal Agency for Support and Evaluation of Graduate Education) for funding this work. We also acknowledge Jéssica Bonomo, Amadeus Azevedo, and Oscar da Silva for technical assistance.

Please cite this article in press as: G. Campani, et al., Recombinant protein production by engineered Escherichia coli in a pressurized airlift bioreactor: A techno-economic analysis, Chem. Eng. Process. (2015), http://dx.doi.org/10.1016/j.cep.2015.10.020

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Please cite this article in press as: G. Campani, et al., Recombinant protein production by engineered Escherichia coli in a pressurized airlift bioreactor: A techno-economic analysis, Chem. Eng. Process. (2015), http://dx.doi.org/10.1016/j.cep.2015.10.020