Studies on growth and biomass formation
Figure 1 showed biomass build up at various batch scales, from shake flask level to fermenters. In shake flasks, the generation of biomass (6.2 mg/ml) was found to be high at 28 h with a maximum specific growth rate (µmax) of 0.20 h-1. However, there was a marked reduction in the duration of log phase with an increase in quantity of biomass was observed when the process was scaled up from lower to higher scale fermenters. Maximum biomass obtained for 19 and 300 L was 11.5 mg/ml (t = 24 h) and 12.8 mg/ml (t = 20 h), respectively. The µmax was also increased and it was shown to be 0.26 h-1 and 0.44 h-1 for 19 and 300 L fermenters respectively. The residual glucose concentration decreased at a high rate in log phase and it was almost depleted in the early stationary phase at the end of 24, 14 and 12 h for flask, 19 and 300 L fermenters, respectively (data not shown) and this showed that the duration of log phase was reduced during scale up. Glucose concentration was nil when the maximum protease activity was observed.
Studies on protease production
The protease production profile during scale up was shown in Fig. 2. For flask cultures, protease secretion was started around 20 h and the maximum activity was obtained around 44 h at stationary phase, with productivity (dp/dt) of 0.48 units h-1 suggesting that the production was not associated in line with biomass generation (Table 1). Similar trend of protease production was also observed for cultures subjected to scale up using fermenters and the maximum protease yield was noticed to be around 20, 24 and 27 units in flask, 19 and 300 L fermenters, respectively. Two distinct phases of growth and production were observed in the process. Though the protease activity was observed to start in cultures drawn in the mid log phase in all scales of operation, the highest activity in the late stationary phase. The protease production was falling in mixed growth associated category as the process did not completely associate with bacterial growth. However, there were reports on protease batches using Bacillus species that revealed that the production was growth associated (Chu et al., 1992; Dumusois and Priest, 1993). In the present study, production was optimal during 24–32 h in fermentors in contrast to 44–48 h in shake flask. Besides, the productivity (dp/dt) was also higher (0.66–1.13 units h-1) using fermenters than shake flask (0.48 units h-1) denoting the advantage that the scale up process is industrially preferable as the process could reduce power, energy and duration with increased operational efficiency. It was reported that increased protease yield with effective oxygen transfer was achieved when the process was scaled up from shake flask to fermentor (Calik et al., 2003).
Table 1
Yield coefficient for biomass and product in shake flask and fermenters
Parameter
|
Flask
|
19 L
|
300 L
|
Yx/s (mg-biomass/mg-glucose)
|
1.69
|
1.80
|
2.21
|
Yp/s (units/mg-glucose)
|
5.78
|
5.82
|
5.64
|
Yp/x (units/mg-biomass)
|
3.42
|
3.24
|
2.56
|
Productivity (dp/dt) (units h− 1)
|
0.48
|
0.66
|
1.13
|
pH and DO profiles
pH variation with a characteristic trend was noticed in all batch processes during scale up (Fig. 3). There was a reduction in the initial pH of 6.9 to 5.4 during growth phase which was followed by an increase in pH above 7.0 during production phase. After complete glucose exhaustion, pH was found to steadily increase during which the protease production was initiated and optimal enzyme production was observed at pH values slightly above neutral. The observation of initial decline followed by rise in pH of the broth could be due to the release of free amino acids during fermentation which was in agreement with an earlier study (Singh et al., 2004). Mixing and aeration efficiency in fermenters enable the availability of high DO content to facilitate bacterial growth followed by product formation with improved process productivity. DO profiles of batches (Fig. 4) showed the following three different phases in fermentation: Phase 1 represented earlier stages of fermentation during which the organism grew profusely. Subsequently, the demand for oxygen increased as evidenced by the drastic reduction of DO content due to high respiration state of the bacterium with the onset of protease production in phase 2. Phase 3 was marked by rise in DO content towards saturation, during which the maximal protease yield was obtained. This characteristic trends were reproduced during scale up though with the reduction in the overall process duration. The protease production was repressed under oxygen limitation as indicated by a sharp decline in the maximum activity when the aeration rate was dropped below 0.8 vvm (Moon and Parulekar, 1991). Hence, through optimal agitation, aeration and foam control operations, the DO content was kept above 20% saturation for supporting the maintenance of active microbial mass to produce the enzyme. Aeration is one of the important process control parameters as most aerobic microbial processes are oxygen limited because of its sparingly soluble nature and maintaining homogeneous DO level throughout the volume of medium in large scale fermenter is challenging (Garcia-Ochoa et al., 2010). Several reports studied the impact of aeration and DO content on protease yield. A study reported the similar trend of DO profile while biomass and protease build up using B.spharicus in both batch and fed batch processes (Singh et al., 2004). Another study reported that there was an increase in protease yield from B. licheniformis when the oxygen supply was reduced (Frankena et al., 1986). It was reported that high biomass and protease yield were achieved at high dissolved oxygen set point in a glucose fed batch fermentation employing Bacillus subtilis culture (Kole et al., 1988).
Mass and oxygen transfer studies
As the performance of aerobic cultures are impacted by oxygen availability, ensuring the mass transfer efficiency of oxygen into liquid is essential and this is done with the measurement of kLa. kLa, OUR, OTR and specific oxygen uptake rate (qO2) were measured for 19 and 300 L at different time points during the fermentation process and the values were presented in Table 2. These values suggested some interesting inferences. During the initial stages of process, at t = 2 h, kLa values were high for both 19 and 300 L fermentors followed by a steady decline till 18 h. The reduction in kLa values was probably due to the increase in biomass which was also marked by a gradual increase in OUR. Reduction in kLa could also be due to increase in the viscosity of fermentation broth, attributed to generation of biomass and other secreted metabolites, led to reduction in the surface area of bubbles caused by the viscous forces generated by fermentation broth (Kilonzo and Margaritis, 2004). The kLa values varied in the range of 0.0049 to 0.0103 s− 1 and 0.0055 to 0.0113 s− 1 for 19 and 300 L fermenters, respectively. It is noteworthy that the values obtained for 19 and 300 L fermenters in different time intervals were rather similar and there was no drastic variation among them. The success of the scale up processes usually relies on maintenance of effective mass transfer rates without much difference in OTR across various scales of operation. During 3–18 h, there was a reduction in the OTR to OUR ratio indicating the highly respiring state of the bacterium. The OUR, which depended on the metabolic functions of the biomass, tend to increase as a result of the increase in the cell formation rate, cell concentration and the substrate consumption rate. After t = 18 h of the process, the OUR decreased because of the decrease in the rates of substrate consumption as well as the cell formation. Around 18 h, OTR/OUR ratio got stabilized during which the secretion of protease was initiated followed by steady increase in OTR to OUR ratio till the end of the process. The steady decrease in OUR during the later period of fermentation revealed that the organism had transformed from a state of actively respiring log phase to relatively low respiring stationary phase wherein the maximum protease yield was obtained. The values for qO2 were found to be in the range of 0.45–23.35 mmol O2 g− 1DCW h− 1 during the process. The peak in the qO2, however, occurred at the inception of the exponential phase and decreased gradually with respect to time. The kLa was specific to a particular fermenter design and also influenced by composition of media and viscosity changes during the fermentation, airflow rate, agitation, batch temperature and the nature of microorganism (Atkinson and Mavituna, 1991; Thiry and Cingolani, 2002; Schmidt, 2005). It was reported earlier that the kLa values for protease production from Bacillus strains varied in the range of 0.01 to 0.055 s− 1 (Calik et al., 2000; 2002; 2003). Another recent study reported that the values for increased protease production by Rhodotorula mucilaginosa ranges from 18 to 135 h− 1 with agitation ranging from 100 to 500 rpm and aeration ranging from 1.0 to 2.5 vvm (Machado et al., 2022).
Table 2
Mass and oxygen transfer properties of culture broth in 19 and 300 L fermenters
Time
(h)
|
kLa
(s− 1)
|
OTR
(mol m− 3s− 1)
|
OUR
(mol m− 3s− 1)
|
OTR/OUR
|
OTRMax
(mol m− 3s− 1)
|
qO2
(mM O2 /g-DCW-h)
|
19 L
|
300 L
|
19 L
|
300 L
|
19 L
|
300 L
|
19 L
|
300 L
|
19 L
|
300 L
|
19 L
|
300 L
|
2
|
0.0103
|
0.0105
|
0.98
|
0.99
|
1.0
|
1.2
|
0.98
|
0.94
|
2.36
|
2.61
|
10.83
|
23.35
|
6
|
0.0059
|
0.0092
|
1.05
|
1.38
|
1.3
|
1.5
|
0.81
|
0.92
|
1.36
|
2.13
|
4.79
|
12.86
|
12
|
0.0052
|
0.0062
|
1.12
|
1.27
|
1.8
|
2.3
|
0.62
|
0.55
|
1.20
|
1.39
|
1.47
|
4.96
|
18
|
0.0049
|
0.0055
|
1.02
|
1.22
|
1.7
|
2.8
|
0.60
|
0.44
|
1.13
|
1.27
|
0.78
|
2.14
|
24
|
0.0066
|
0.0071
|
1.17
|
1.56
|
1.0
|
2.3
|
1.17
|
0.68
|
1.52
|
1.64
|
0.58
|
1.34
|
30
|
0.0079
|
0.0126
|
1.24
|
2.51
|
0.7
|
1.2
|
1.77
|
2.09
|
1.82
|
2.91
|
0.45
|
0.72
|
The DO concentration in the culture broth depends on the OTR from the gas to the liquid phase and also on the rate of its consumption by the microorganism, OUR. Though OUR is consecutive to OTR, but, the consumption of oxygen can affect OTR, and hence, the oxygen dynamics has to be assessed and optimized with respect to changing microbial dynamics and culture viability. Though OUR measurement has received the due attention in various bioprocesses, the reports on its significance in protease fermentation are limited (Garcia-Ochoa et al., 2010).
Oxygen tension and shear forces due to aeration and agitation affect the mass transfer properties of the medium during fermentation. The effect of aeration and agitation on kLa was also studied using the broth at their respective conditions. The kLa for various agitation states such as 100, 200, 300, 400 and 500 rpm at aeration rate of 1 vvm was found to be 0.009, 0.014, 0.021, 0.029 and 0.036 respectively showing that increased agitation resulted in increased mass transfer features; however, generation of biomass was adversely affected by high agitation rates due to acting of shear forces on cells that reduced both biomass and protease yield. As fermentation progressed, there is a need to increase the agitation rate for promoting bacterial biomass to consume the substrate. At low agitation rates especially when the viscosity of medium increases during fermentation, the substrate uptake is usually low and this is mainly due to formation of dead-zones, which supported the mass transfer resistance. Hence, the kLa value of 0.029 s− 1 with agitation speed of 400 rpm at aeration of 1 vvm was found to be optimal for protease production in this study. The agitation in the range of 150–300 rpm was reported to be optimal for protease production using Bacillus group of bacteria (Potumarthi et al., 2007; Bhunia et al., 2012). The kLa values for protease bioprocess using Bacillus licheniformis NCIM-2042 was reported to be 41, 54 and 56 at 1, 2 and 3 vvm of airflow rate respectively; but, the maximum protease yield was obtained at 2 vvm and with low agitation speed of 180 rpm (Dey et al., 2016). On the other hand, it was also reported that low agitation affected the yield of protease by Bacillus sp. due to inadequate oxygen transfer (Calik et al., 2000; Joo et al., 2002; Genckal and Tari, 2006; Machado et al., 2022; Bhunia et al., 2012). Agitation is reported to be more effective than aeration for enzyme production as high agitation generates bubbles that increase the gas–liquid interface area as well as the residence time in the culture medium causing a higher rate of DO (Teruasmaki et al., 2018). Shear stress due to oxygen supply through high aeration rate adversely affect the organism’s morphology, physiology, and consequently, the biomass generation and enzyme production (Machado et al., 2022).
Besides studies on mass transfer properties, other scale up parameters such as impeller tip speed, aeration rate, superficial gas velocity, and gas hold up that were influenced by aeration and agitation were also evaluated (Table 3). The results showed that the values obtained for 300 L fermenter were higher than 19 L fermenter substantiating the observation that effective aeration and mixing were attained in 300 L fermenter that resulted in improved productivity of the process when compared to 19 L. The values of high gas hold up and impeller tip speed in 300 L fermenter indicated reduction in bubble size or increase in the number of bubbles which, in turn, enhanced mass transfer area and thus the efficacy of the process when performed in high scales of operation. In viscous broth, the collisions among bubbles lead to the formation of large bubbles that reduce the mass transfer interfacial area (Tirunehe and Norddahl, 2016). The stirrer speed and the mixing intensity play a major role in the breaking up of bubbles in viscous broth (Garcia-Ochoa and Gomez, 2009). With an increase in viscosity of broth, achieving turbulent flow was possible by increasing stirrer speed and the mixing intensity.
Table 3
Scale up parameters for the culture broth in 19 and 300 L fermenters
Parameter
|
19 L
|
300 L
|
Impeller tip speed (m/s)
|
3.768
|
5.233
|
Aeration rate (m3/m3s)
|
12.45 x 10− 3
|
7.4 x 10− 3
|
Superficial gas velocity (m/s)
|
3.134 x 10− 3
|
8.172 x 10− 3
|
Gas hold up (m3s/m3)
|
80.32
|
135.14
|
Impeller tip speed = 2 π n di |
(di = diameter of impeller = 0.080 m for 19L and 0.200 m for 300L; n = 450/60 for 19 L and 250/60 for 300 L fermenter) |
Aeration rate, Ar = Fg/Vr |
Fg = volumetric gas flow rate, Vr = volume of the reactor (19 L) |
Superficial gas velocity = Fg/A |
A = Area of cross section |
Gas hold up = Vr/Fg= 1/Ar |
Impact of culture rheology in fermentation scale up
The production of protease is largely influenced by OTR that in turn depends on medium rheology. The effect of rheological features of the medium during scale up of enzyme production was investigated by analyzing viscosity and density of the broth in 19 and 300 L fermentors, at various time intervals. It was observed that the viscosity of the broth in both fermentors during cultivation was found to be increasing (t = 0 to 36 h) due to formation of biomass as well as accumulation of metabolites indicating the Non-Newtonian rheological features of the broth (Table 4). Rheological behavior the production medium broth with respect to time is of paramount importance in understanding the gas and nutrient transport events in the fermenter. Effective mass transfer through mixing of medium is achieved by means of either reduction in medium viscosity or higher turbulent flow in the culture (Genekoplis, 1999). Reduction in the medium viscosity was not feasible in the current study due to changing dynamics of the medium during the course of fermentation and the rheological dynamics is attributed to the nature of microbial growth, metabolites and medium composition. However, effective mass transfer and mixing of medium could be ensured by higher turbulent flow by providing optimal stirrer speed (N). The values of Reynold’s number (RNe) for the culture broth showed that the turbulent flow (RNe >104) was achieved in both fermenter systems. However, the values were almost three fold higher for the culture developed in 300 L fermenter than 19 L. This performance efficacy in large scale fermentation system could be attributed to vessel geometrics, optimal agitation, high surface area of air bubbles, oxygen diffusivity, effective mass transfer and mixing features (Kilonzo and Margaritis, 2004). However, it was reported that challenges associated with poor bulk mixing and gas–liquid mass transfer were less in smaller fermenters than that of the large scale operations (Enfors et al., 2001; Schmidt, 2005). Nevertheless the culture displayed a marginal variation in rheological features when subjected to different scales of process, effective mixing through higher turbulent flow of broth in 300 L resulted in increased productivity. Hence, optimal stirrer speed (N) of 450 and 250 rpm was required to achieve both turbulent flow and higher protease yield in 19 and 300 L fermentors, respectively. High protein concentration in the medium not only increase the viscosity of the culture accompanied by reduction in oxygen transfer rates but also overall duration of the operation that affected the batch productivity (Potumarthi et al., 2007).
Table 4
Rheological features of culture broth during protease production in fermenters
Parameter
|
19 L
|
300 L
|
0 h
|
12 h
|
24 h
|
36 h
|
0 h
|
12 h
|
24 h
|
36 h
|
Density (ρmedium)
(Kg/m3)
|
1024
|
1083
|
1124
|
1148
|
1025
|
1119
|
1138
|
1172
|
Viscosity (µsol)
(Kg/m2s)
|
8.5 x 10− 4
|
9.44 x 10− 4
|
10.27 x 10− 4
|
10.73 x 10− 4
|
8.61 x 10− 4
|
10.11 x 10− 4
|
11.47 x 10− 4
|
12.06 x 10− 4
|
Reynold’s Number
RNe= (ND2ρ/µ)
|
5.78 x 104
|
5.51 x 104
|
5.25 x 104
|
5.14 x 104
|
19.9 x 104
|
18.5 x 104
|
16.54 x 104
|
16.19 x 104
|
a) ρ medium = (Weight of medium/ Weight of H2O at 300 C) x ρ water at 300 C |
Weights of the fermentation medium in 19 L and 300 L fermenters for 0, 12, 24 and 36 h are 102.4, 108.3, 112.4 and 114.8g and 102.5, 111.9, 113.8 and 115.9g respectively. Weight of H2O at 300 C is 100g. |
ρ water at 300 C is 1g/cm3 |
b) µsol = [( ρ medium x time flow of medium)/( ρwater x time flow of water)] x µwater at 300C |
Time flow of medium for 0, 12, 24 and 36 h of 19 and 300 L are 79, 83, 87 and 89g and 80, 86, 96 and 98g respectively. |
Time flow of water is 76 s. µwater at 300C is 7.9818 x10− 4 Kg/m2s |
c) RNe = (ND2 ρ/µ) |
D - Diameter of impeller of fermentor and it is 80 and 200 x 10− 3 m for 19 and 300 L fermenters, respectively. |
N - Stirrer speed and is 450/60 and 250/60 (revolutions per sec) for 19 and 300 L fermenters, respectively. |
For Turbulent flow, RNe > 10000. For Linear flow, RNe < 10000 |