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Hydrodynamic and mass transfer characterization of a down flow jet loop bioreactor
Biochemical Engineering Journal 8 (2001) 241–250
Hydrodynamic and mass transfer characterization of a down flow jet loop bioreactor Amir Masoud Jamshidi, Morteza Sohrabi∗ , Farzaneh Vahabzadeh, Babak Bonakdarpour Department of Chemical Engineering, Amirkabir University of Technology, Tehran 15, Iran Received 16 February 2000; accepted 19 March 2001
Abstract The effects of certain pertinent parameters such as gas and liquid flow rates and nozzle position on the behavior of a down flow jet loop bioreactor (DJB) have been studied. The mean residence times of gas and liquid phases and the gas hold-up within the bioreactor have been measured. A correlation has been presented for the gas hold-up. In addition, the overall volumetric mass transfer coefficient, and the influence of the gas flow rate and the position of the nozzle inside the draft tube on the latter has been determined and a correlation has been derived. The liquid residence time distribution (RTD) within the bioreactor has been studied by tracer analysis and the mixing time was calculated using the RTD results. © 2001 Elsevier Science B.V. All rights reserved. Keywords: Jet loop bioreactor; Down flow bioreactor; Volumetric mass transfer coefficient; Residence time distribution; Gas hold-up
1. Introduction Jet loop bioreactors are being widely applied in a number of chemical and biochemical processes which involve gas–liquid reactions. For example, recently such bioreactors have been employed in studies on the biodesulfurization of petroleum [1]. This is due to certain specific properties of these bioreactors such as high mass transfer performance, good dispersing effects and relatively low power requirements. In a jet loop bioreactor circulation and gas dispersion is achieved by a liquid jet drive. Liquid and gas are injected into the bioreactor with high velocity, via a spray nozzle located inside the draft tube [2,3]. In Fig. 1, the manner in which the gas and liquid move inside in the DJB is shown. As seen, the gas enters the bioreactor at the position in which the spray nozzle is located inside the draft tube. It then crosses the whole length of the draft tube below the nozzle in the downward direction after which it enters the annulus and flows upwards and exits from the top of the bioreactor. The direction of the liquid flow is also shown in Fig. 1. Compared to the bioreactors in which the spray nozzle is placed at the bottom of the draft tube, this arrangement has the advantage that the gas has a higher residence time inside the bioreactor and a move pronounced and homogeneous mixing is achieved inside the bioreactor. The increased gas resi∗
Corresponding author.
dence time will result in a higher mass transfer which makes such bioreactors suitable for conducting either aerobic fermentation or anaerobic fermentations in which a gas phase such as hydrogen needs to be introduced into the fermentor. Compare to a air lift fermentors, DJBs have the advantage that the extent of mixing and the distribution of bubble sizes can be more directly controlled. The high extent of mixing that can be achieved in such bioreactors is advantageous in three phase system such as fermentations in which an organic phase such as petroleum is employed as the substance. The most important parameter in multiphase systems is the generation of interfacial area, the extent of which is directly related to the gas hold-up. The gas hold-up in turn depends upon the physical properties of the liquid, the flow regime and bioreactor efficiency [4]. The DJBs in comparison with other types of gas–liquid bioreactors provide higher surface area between the two phases. Design of appropriate nozzles and location of the latter in the draft tube are extremely important with regards to the extent of gas dispersion in liquid phase [5]. Fixing the nozzle at the higher sections of the draft tube enables the DJB to be applied in processes involving suspended solid particles [6]. As a typical example of such processes waste water treatment may be mentioned [7]. In the present study, the hydrodynamic and mass transfer behavior of DJB at various gas and liquid flow rates as well as various location of the spray nozzle has been investigated.
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Nomenclature A AD B C∗ CL Dd Deq DG Di DL Do Dr (E/V)L Fr g HL HLG Hn Hp Hr KL a Ld Mo QG QL Rej s Sc Sh t UL USG V VL Greek ε µL ρL σL τG τL
cross-sectional area of the bioreactor (m2 ) flow area (π /4(Do2 − Di2 )) (m2 ) 1/KL a saturation oxygen concentration in liquid (mg/l) oxygen concentration in liquid (mg/l) draft tube diameter (m) equivalent flow diameter (Do2 − Di2 )0.5 (m) air hole diameter in nozzle (m) smaller diameter in liquid annular hole in the nozzle (m) diffusion coefficient (m2 /s) bigger diameter in liquid annular hole in the nozzle (m) bioreactor column diameter (m) energy dissipation rate per unit volume ((ρL × AD × UL3 )/2V ) (W/m3 ) Froud number (USG (gDr )−0.5 ) gravitational constant (m/s2 ) level of the clear liquid (cm) level of the two phase dispersion (cm) distance between the upper edge of the draft tube and the nozzle (cm) distance between in lower edge of the draft tube and the impact plate (m) bioreactor height (m) volumetric mass transfer coefficient (s−1 ) height of the draft tube (cm) Morton number (µ4L g/σL3 ρL ) gas flow rate (m3 /s) liquid flow rate (m3 /s) jet Reynolds number ((D eq × U L × ρ L )/µL ) electrode time lag (s) Schmidt number (µL /ρ L DL ) Sherwood number (KL a × Dr2 /DL ) time (s) linear liquid velocity based on AD (m/s) superficial gas velocity based on A (m/s) two phase volume in the bioreactor (m3 ) liquid volume in the bioreactor (m3 ) symbols gas hold-up liquid viscosity (MPa s) liquid density (kg/m3 ) liquid surface tension (N/m) gas phase mean residence time (s) liquid phase mean residence time (s)
Fig. 1. Downflow jet loop bioreactor with coaxial draft tube.
2. Materials and methods The experimental setup is shown in Fig. 2. The bioreactor consisted of a vertical glass tube, 0.137 m in diameter and 0.5 m in height. A draft tube was located axially in the center of the bioreactor and fixed at a short distance from the base. The spray nozzle was placed inside the draft tube. Location of the spray nozzle could be varied from the top to the bottom of the draft tube. Other dimensions of the experimental apparatus are given in Table 1. The design of the spray nozzle is shown in Fig. 3. According to Fig. 3 the nozzle consisted of two main sections — the central part and the main body. There was no contact between the two phases inside the nozzle. The gas, after leaving the compressor, passed a gas flow meter and entered the nozzle via a central port, 1 mm diameter. The liquid phase entered the inside the bioreactor nozzle via an annular opening, 6 mm i.d. and 8 mm o.d. The liquid was withdrawn below the impact plate and recycled to the nozzle via a flow meter by means of a pump. 3. Results and discussion 3.1. Bioreactor hydrodynamics 3.1.1. Mean residence time and gas hold-up in the bioreactor The liquid phase mean residence time could be calculated from the following equation: τL =
VL liquid hold up = liquid flow rate QL
(1)
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Fig. 3. Cross-section of the spray nozzle.
Fig. 2. The experimental set-up: 1, reactor; 2, draft tube; 3, spray nozzle; 4, nozzle support; 5, gas outlet; 6, liquid flow meter; 7, liquid pressure gage; 8, gas pressure gage; 9, gas flow meter; 10, camera; 11, impact plate; 12, pump; 13, rotary sample; 14, oxygen meter; 15, compressor; 16, pulse injection system.
The liquid flow rate was varied from 70 to 160 cm3 /s and the liquid hold up within the bioreactor was 4750 cm3 . Hence, the range of liquid mean residence times was between 30 and 68 s. The gas hold-up, ε, in the bioreactor during steady-state was measured from the difference between the level of the two phase dispersion, HLG , and that of the clear liquid HL (the gas flow rate, QG , was determined from the gas flow meter and gas loop pressure gauge). ε=
HLG − HL HLG
(2)
Gas hold-up was measured by the volumetric expansion method and hydrostatic pressure determination. Depending upon the flow rates of gas and liquid, gas hold-up varied from 0.01 to 0.204. Gas phase mean residence time was calculated using the following equation: τG =
VL ε QG 1 − ε
(3)
Regarding the operating conditions, the mean residence time of gas phase was found to be between 0.6 and 1.5 s. In Fig. 4, the effect of nozzle position inside the draft tube on gas hold-up is shown. It may be observed from this figure that the gas hold-up is increased by increase in gas flow rate and decrease in the immersion height of the nozzle inside the draft tube. The total volume of bioreactor crossed by the gas increases when the nozzle is raised inside the draft
Table 1 Dimensions of down flow jet loop bioreactor Bioreactor height (m) Bioreactor diameter (m) Draft tube diameter (m) Ratio of draft tube diameter to bioreactor diameter Draft tube length (m) Nozzle immersion height (m) Distance between lower edge of the draft tube and the impact plate (m) Air hole diameter in nozzle (m) Liquid annular hole inside diameter in nozzle (m) Liquid annular hole outside diameter in nozzle (m) Clean liquid level height in bioreactor (m) Two phase dispersion level height in the bioreactor (m)
Hr Dr Dd Dd /Dr Ld Hn Hp DG Di Do HL HLG
0.5 0.137 0.051 0.372 0.26 0, 0.115, 0.25 0.045 0.001 0.006 0.008 0.35 0.35–0.45
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Fig. 5. Influence of liquid flow rate on gas hold-up.
Fig. 4. Influence of nozzle position on gas hold-up.
tube. However, it should be pointed out that this only results in an increase in gas hold-up if the gas stream actually hits the impact plate. If the nozzle is raised inside the draft tube such that the gas stream does not hit the impact plate, for example at very high gas flow rates, then increase in gas hold-up will not be expected since collision with the impact plate leads to smaller bubble sizes which results in higher value of ε inside the bioreactor. In Figs. 5–7, the effects of liquid and gas flow rates on the gas hold-up are shown. Each figure corresponds to a given position of nozzle in the draft tube. It is clear from these figures that for a given position of the nozzle, increase in liquid or gas flow rates promote gas hold-up. Variations of gas hold-up as a function of liquid energy dissipation rate per unit volume of the bioreactor (E/V)L are given in Figs. 8–10. Velan and Ramanujam [3] presented a correlation for gas hold-up inside down flow jet loop reactors based on dimensionless parameters. Jamshidi [1] modified this correlation to include the effect of draft tube height and the position of the nozzle inside the draft tube as follows: ε = 1.1072 × 10−4 (Rej )1.095 (Fr)0.885
1.4Ld 1 + 2Hn
0.9
Dd Dr
0.025
(4)
Fig. 6. Influence of liquid flow rate on gas hold-up.
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Fig. 9. Influence of energy dissipation rate (E/V)L on gas hold-up. Fig. 7. Influence of liquid flow rate on gas hold-up.
Fig. 8. Influence of energy dissipation rate (E/V)L on gas hold-up.
Fig. 10. Influence of energy dissipation rate (E/V)L on gas hold-up.
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Fig. 11. Experimental and predicted gas hold-up in air–water system.
In Fig. 11, a comparison has been made between the gas hold-up experimental data and those determined using Eq. (4). The relative error is ±20%. In the above equation four dimensionless terms have been employed. The first two terms reflect the operating conditions inside the bioreactor and have a very important significance. Rej reflects the liquid flow situation in the nozzle exit port. Given the power to which Rej is raised in the above equation it can be seen that hold-up varies almost linearly with the liquid linear velocity UL , therefore, doubling of the liquid flow rate will results in a more than double increase in the gas hold-up. The second dimensionless number in the above correlation is the Froud number which has a direct relationship with gas superficial velocity. Therefore, according to the above empirical equation gas hold-up is related to the gas flow rate raised to the power 0.885. In the other words a doubling of the gas flow rate will result in a 85% increase in the gas hold-up. This means that the effect of the increase in liquid flow rate on the gas hold-up is more important than the corresponding effect of the increase of the gas flow rate. 3.1.2. Volumetric mass transfer coefficient To determine the mass transfer coefficient between the two phases, the air–water system was used. Measurement of dissolved oxygen in the fluid was performed using an oxygen meter (OXI 325-A/Set Best No. 200216 wtw) connected to an electrode. A batch of deionized water was deaerated by
bubbling argon gas. The experiment was started by passing deaerated water and air through the nozzle. Dissolved oxygen concentration was measured at intervals of 5 s. In each experiment, about 50 readings were recorded. Experiments were performed at various gas flow rates and at different locations of nozzle in the draft tube and a constant temperature of 25 ± 0.1◦ C. Oxygen saturation in water was achieved within 40–140 s, depending upon the hydrodynamic conditions. Volumetric mass transfer coefficients were calculated from the following equation: dCL = KL a(C ∗ − CL ) dt
(5)
where CL is the dissolved oxygen concentration at any time, t, C∗ is the saturation concentration of oxygen in liquid and KL a is the volumetric mass transfer coefficient. Integration of the above equation with the following boundary conditions: t = 0,
CL = Co
t = t,
CL = CL
yields the following relation: ln
C ∗ − Co = KL a t C ∗ − CL
(6)
Eq. (6) is applicable only if the flow of fluid within the bioreactor is fully back mixed. In the transient technique, the
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dynamic conditions of the oxygen electrode affect the KL a measurement. Normally, a delay or time lag is encountered in electrode response. However, the following first-order model was applied to compensate for the delay [8–10]. −t/B C ∗ − CL e−t/s Bs e − (7) = C ∗ − Co s B B −s where s and B are electrode time lag and that of 1/KL a, respectively. If t s, then Eq. (5) may be simplified as C ∗ − CL B e−t/B = C ∗ − Co B −s
(8)
Thus, the slope and intercept of the plot of ln(C ∗ − CL )/(C ∗ − Co ) versus time, gives −1/K L a = −B and ln(B/B − s), respectively; from the latter the electrode delay could be calculated. In Fig. 12, the effect of gas flow rate on the volumetric mass transfer coefficient for various locations of nozzle is shown. It is clear that increase in gas flow rate increases KL a. In addition, the position of nozzle inside the draft tube has a marked effect on KL a (thus, increasing the height of the nozzle position from the bottom of the draft tube, increases KL a). Velan and Ramanujam [4] presented a correlation for KL a inside a down flow jet loop reactors based on dimensionless parameters. Jamshidi [1] modified this correlation to include the effect of draft tube height and the position of the sprayer
Fig. 12. Influence of gas flow rate on volumetric mass transfer coefficient.
Fig. 13. Experimental and predicted mass transfer coefficient.
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inside the draft tube and came up with following correlation: Sh = 1.76(Sc)0.5 (Rej )0.99 (Fr)0.53 (Mo)−0.06 0.15 + 1.05Ld 0.71 × Hn + L d
(9)
In Fig. 13, a comparison has been made between the experimental KL a data and those estimated using Eq. (9). The degree of agreement was within ±11%. The above equation contains five dimensionless terms. The dimensionless number Rej , Fr and Mo reflects the hydrodynamic condition inside the bioreactor whilst Sc reflects the mass transfer condition inside the bioreactor. The last dimensionless term in the above equation shows the effect of the position of the nozzle in the draft tube. The exponent on Rej and Fr indicate the importance of these parameters with regards to the overall volumetric oxygen transfer coefficient. Comparison of the (4) and (9) shows
the effect of the variation of Rej and the linear liquid velocity on the overall volumetric oxygen transfer coefficient is similar to the effect of variation of these dimensionless parameters on gas hold-up whilst the effect of the variation of Fr on the volumetric oxygen transfer coefficient is appreciably less than the effect of this parameter on gas hold-up. In other words, the effect of the increase in the liquid flow rate on gas hold-up is noticeably higher than the effect gas flow rate on this parameter. This shows that in the calculation of both gas hold-up and KL a, hydrodynamic parameters play a significant role. 3.1.3. Residence time distribution (RTD) of liquid in the bioreactor To determine the RTD of liquid within the bioreactor, tracer analyses were performed. It was evident that the liquid was circulated in a closed loop through the system, while the flow of gas was continuous. It was anticipated, therefore, that a pulse of a tracer injected in the liquid flow at the entrance of the bioreactor will be evenly distributed throughout the
Fig. 14. Response at the exit of a down flow jet loop reactor subsequent to an impulse input.
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Fig. 15. Response at the exit of a down flow jet loop reactor subsequent to an impulse input.
system after a short period of time. The concentration of tracer at the bioreactor outlet might be, therefore, regarded as a measure of the hydrodynamic conditions and the pattern of flow within the bioreactor [11]. In each experiment, 1 cm3 of a known solution of potassium permanganate was injected rapidly via the liquid flow inlet to the bioreactor. Samples from the liquid phase at the outlet of the bioreactor were collected, using a circular vessel divided into 23 segments, rotating at desired speeds. By this method, successive samples were obtained at equal time intervals. The accurate determination of angular speed of the vessel was achieved using the digital Photo/Contact Tachometer DT 2236. The optical densities of samples were measured by a Pye Unicam SP 1700 UV Spectrophotometer. The relative concentrations (C/Cmax ) of samples were then calculated and RTD curves were plotted.
The tracer analysis were conducted at various liquid flow rate (5000–9000 cm3 /min) and gas flow rates (3000–50000 cm3 /min). In addition the distance between the nozzle and the top of the draft tube was varied in the range of 0–25 cm. In Figs. 14 and 15, the relative concentration (C/Cmax ) of tracer at the bioreactor outlet have been plotted versus time. Tracer analysis under different experimental conditions (various gas and liquid flow rates, position of nozzle) was performed. However, the overall patterns of RTD curves obtained remained almost unaffected. Mixing time within the bioreactor was calculated using the RTD results. In Fig. 16, the effects of inlet gas velocity and position of nozzle on mixing time are observed. It is evident that the increase in gas velocity promotes the mixing process, and hence, decreases the mixing time. Locating the nozzle at higher levels of the draft tube would enhance the
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Fig. 16. Effect of superficial gas velocity on mixing time.
circulation within the bioreactor and consequently decrease the mixing time.
4. Conclusion The following conclusions may be drawn from the present results: 1. The gas hold-up increases with increase in gas and liquid flow rates and the height of nozzle position from the bottom of the draft tube. However, the effect of gas flow rate is more pronounced. 2. The volumetric mass transfer coefficient increases with increases in gas and liquid flow rates and the height of the nozzle position from the bottom of the draft tube. 3. It may be observed from the RTD curves that the tracer concentration within the bioreactor, attains a uniform concentration after 2–6 s. Absence of long tails in the RTD curves indicates that dead regions are not present within the bioreactor. 4. The mixing time is a function of gas and liquid flow rates and the nozzle position in the draft tube. Increasing the
fluid flow rates and the height of nozzle location in the draft tube, reduces the mixing time. References [1] A.M. Jamshidi, Chemical engineering, Ph.D. Thesis, Amirkabir University of Technology, Tehran, Iran (2000). [2] C.A.M.C. Dirix, K. Van der wiele, Chem. Eng. Sci. 45 (8) (1990) 2333–2340. [3] M. Velan, T.K. Ramanujam, Chem. Eng. Sci. 47 (9–11) (1992) 2871–2876. [4] M. Velan, T.K. Ramanujam, Can. J. Chem. Eng. 69 (1991) 1257–1261. [5] N. Rabiger, A. Vogelpohl, Chem. Ing. Tech. 55 (6) (1983) 486–487. [6] Karl Schugerl, Bioreaction Engineering, Vol. 2, Wiley, New York, 1991. [7] K. Kulkarni, Y. Shah, A. Schumpe, Chem. Eng. Commun. 24 (1983) 307–337. [8] M. Nakanoh, F. Yoshida, Ind. Eng. Chem. Process Des. Dev. 19 (1980) 190. [9] S.A. El-Temtamy, S.A. Khalil, A.A. Nour-El-Din, A. Gaber, Appl. Microbiol. Biotechnol. 19 (1984) 376. [10] M.Y. Chisti, M. Moo-Young, Biotech. Bioeng. (1988) 487–494. [11] M. Sohrabi, A.M. Jamshidi, J. Chem. Tech. Biotechnol. 69 (1997) 415–420.
Hydrodynamic and mass transfer characterization of a down flow jet loop bioreactor Amir Masoud Jamshidi, Morteza Sohrabi∗ , Farzaneh Vahabzadeh, Babak Bonakdarpour Department of Chemical Engineering, Amirkabir University of Technology, Tehran 15, Iran Received 16 February 2000; accepted 19 March 2001
Abstract The effects of certain pertinent parameters such as gas and liquid flow rates and nozzle position on the behavior of a down flow jet loop bioreactor (DJB) have been studied. The mean residence times of gas and liquid phases and the gas hold-up within the bioreactor have been measured. A correlation has been presented for the gas hold-up. In addition, the overall volumetric mass transfer coefficient, and the influence of the gas flow rate and the position of the nozzle inside the draft tube on the latter has been determined and a correlation has been derived. The liquid residence time distribution (RTD) within the bioreactor has been studied by tracer analysis and the mixing time was calculated using the RTD results. © 2001 Elsevier Science B.V. All rights reserved. Keywords: Jet loop bioreactor; Down flow bioreactor; Volumetric mass transfer coefficient; Residence time distribution; Gas hold-up
1. Introduction Jet loop bioreactors are being widely applied in a number of chemical and biochemical processes which involve gas–liquid reactions. For example, recently such bioreactors have been employed in studies on the biodesulfurization of petroleum [1]. This is due to certain specific properties of these bioreactors such as high mass transfer performance, good dispersing effects and relatively low power requirements. In a jet loop bioreactor circulation and gas dispersion is achieved by a liquid jet drive. Liquid and gas are injected into the bioreactor with high velocity, via a spray nozzle located inside the draft tube [2,3]. In Fig. 1, the manner in which the gas and liquid move inside in the DJB is shown. As seen, the gas enters the bioreactor at the position in which the spray nozzle is located inside the draft tube. It then crosses the whole length of the draft tube below the nozzle in the downward direction after which it enters the annulus and flows upwards and exits from the top of the bioreactor. The direction of the liquid flow is also shown in Fig. 1. Compared to the bioreactors in which the spray nozzle is placed at the bottom of the draft tube, this arrangement has the advantage that the gas has a higher residence time inside the bioreactor and a move pronounced and homogeneous mixing is achieved inside the bioreactor. The increased gas resi∗
Corresponding author.
dence time will result in a higher mass transfer which makes such bioreactors suitable for conducting either aerobic fermentation or anaerobic fermentations in which a gas phase such as hydrogen needs to be introduced into the fermentor. Compare to a air lift fermentors, DJBs have the advantage that the extent of mixing and the distribution of bubble sizes can be more directly controlled. The high extent of mixing that can be achieved in such bioreactors is advantageous in three phase system such as fermentations in which an organic phase such as petroleum is employed as the substance. The most important parameter in multiphase systems is the generation of interfacial area, the extent of which is directly related to the gas hold-up. The gas hold-up in turn depends upon the physical properties of the liquid, the flow regime and bioreactor efficiency [4]. The DJBs in comparison with other types of gas–liquid bioreactors provide higher surface area between the two phases. Design of appropriate nozzles and location of the latter in the draft tube are extremely important with regards to the extent of gas dispersion in liquid phase [5]. Fixing the nozzle at the higher sections of the draft tube enables the DJB to be applied in processes involving suspended solid particles [6]. As a typical example of such processes waste water treatment may be mentioned [7]. In the present study, the hydrodynamic and mass transfer behavior of DJB at various gas and liquid flow rates as well as various location of the spray nozzle has been investigated.
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Nomenclature A AD B C∗ CL Dd Deq DG Di DL Do Dr (E/V)L Fr g HL HLG Hn Hp Hr KL a Ld Mo QG QL Rej s Sc Sh t UL USG V VL Greek ε µL ρL σL τG τL
cross-sectional area of the bioreactor (m2 ) flow area (π /4(Do2 − Di2 )) (m2 ) 1/KL a saturation oxygen concentration in liquid (mg/l) oxygen concentration in liquid (mg/l) draft tube diameter (m) equivalent flow diameter (Do2 − Di2 )0.5 (m) air hole diameter in nozzle (m) smaller diameter in liquid annular hole in the nozzle (m) diffusion coefficient (m2 /s) bigger diameter in liquid annular hole in the nozzle (m) bioreactor column diameter (m) energy dissipation rate per unit volume ((ρL × AD × UL3 )/2V ) (W/m3 ) Froud number (USG (gDr )−0.5 ) gravitational constant (m/s2 ) level of the clear liquid (cm) level of the two phase dispersion (cm) distance between the upper edge of the draft tube and the nozzle (cm) distance between in lower edge of the draft tube and the impact plate (m) bioreactor height (m) volumetric mass transfer coefficient (s−1 ) height of the draft tube (cm) Morton number (µ4L g/σL3 ρL ) gas flow rate (m3 /s) liquid flow rate (m3 /s) jet Reynolds number ((D eq × U L × ρ L )/µL ) electrode time lag (s) Schmidt number (µL /ρ L DL ) Sherwood number (KL a × Dr2 /DL ) time (s) linear liquid velocity based on AD (m/s) superficial gas velocity based on A (m/s) two phase volume in the bioreactor (m3 ) liquid volume in the bioreactor (m3 ) symbols gas hold-up liquid viscosity (MPa s) liquid density (kg/m3 ) liquid surface tension (N/m) gas phase mean residence time (s) liquid phase mean residence time (s)
Fig. 1. Downflow jet loop bioreactor with coaxial draft tube.
2. Materials and methods The experimental setup is shown in Fig. 2. The bioreactor consisted of a vertical glass tube, 0.137 m in diameter and 0.5 m in height. A draft tube was located axially in the center of the bioreactor and fixed at a short distance from the base. The spray nozzle was placed inside the draft tube. Location of the spray nozzle could be varied from the top to the bottom of the draft tube. Other dimensions of the experimental apparatus are given in Table 1. The design of the spray nozzle is shown in Fig. 3. According to Fig. 3 the nozzle consisted of two main sections — the central part and the main body. There was no contact between the two phases inside the nozzle. The gas, after leaving the compressor, passed a gas flow meter and entered the nozzle via a central port, 1 mm diameter. The liquid phase entered the inside the bioreactor nozzle via an annular opening, 6 mm i.d. and 8 mm o.d. The liquid was withdrawn below the impact plate and recycled to the nozzle via a flow meter by means of a pump. 3. Results and discussion 3.1. Bioreactor hydrodynamics 3.1.1. Mean residence time and gas hold-up in the bioreactor The liquid phase mean residence time could be calculated from the following equation: τL =
VL liquid hold up = liquid flow rate QL
(1)
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Fig. 3. Cross-section of the spray nozzle.
Fig. 2. The experimental set-up: 1, reactor; 2, draft tube; 3, spray nozzle; 4, nozzle support; 5, gas outlet; 6, liquid flow meter; 7, liquid pressure gage; 8, gas pressure gage; 9, gas flow meter; 10, camera; 11, impact plate; 12, pump; 13, rotary sample; 14, oxygen meter; 15, compressor; 16, pulse injection system.
The liquid flow rate was varied from 70 to 160 cm3 /s and the liquid hold up within the bioreactor was 4750 cm3 . Hence, the range of liquid mean residence times was between 30 and 68 s. The gas hold-up, ε, in the bioreactor during steady-state was measured from the difference between the level of the two phase dispersion, HLG , and that of the clear liquid HL (the gas flow rate, QG , was determined from the gas flow meter and gas loop pressure gauge). ε=
HLG − HL HLG
(2)
Gas hold-up was measured by the volumetric expansion method and hydrostatic pressure determination. Depending upon the flow rates of gas and liquid, gas hold-up varied from 0.01 to 0.204. Gas phase mean residence time was calculated using the following equation: τG =
VL ε QG 1 − ε
(3)
Regarding the operating conditions, the mean residence time of gas phase was found to be between 0.6 and 1.5 s. In Fig. 4, the effect of nozzle position inside the draft tube on gas hold-up is shown. It may be observed from this figure that the gas hold-up is increased by increase in gas flow rate and decrease in the immersion height of the nozzle inside the draft tube. The total volume of bioreactor crossed by the gas increases when the nozzle is raised inside the draft
Table 1 Dimensions of down flow jet loop bioreactor Bioreactor height (m) Bioreactor diameter (m) Draft tube diameter (m) Ratio of draft tube diameter to bioreactor diameter Draft tube length (m) Nozzle immersion height (m) Distance between lower edge of the draft tube and the impact plate (m) Air hole diameter in nozzle (m) Liquid annular hole inside diameter in nozzle (m) Liquid annular hole outside diameter in nozzle (m) Clean liquid level height in bioreactor (m) Two phase dispersion level height in the bioreactor (m)
Hr Dr Dd Dd /Dr Ld Hn Hp DG Di Do HL HLG
0.5 0.137 0.051 0.372 0.26 0, 0.115, 0.25 0.045 0.001 0.006 0.008 0.35 0.35–0.45
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Fig. 5. Influence of liquid flow rate on gas hold-up.
Fig. 4. Influence of nozzle position on gas hold-up.
tube. However, it should be pointed out that this only results in an increase in gas hold-up if the gas stream actually hits the impact plate. If the nozzle is raised inside the draft tube such that the gas stream does not hit the impact plate, for example at very high gas flow rates, then increase in gas hold-up will not be expected since collision with the impact plate leads to smaller bubble sizes which results in higher value of ε inside the bioreactor. In Figs. 5–7, the effects of liquid and gas flow rates on the gas hold-up are shown. Each figure corresponds to a given position of nozzle in the draft tube. It is clear from these figures that for a given position of the nozzle, increase in liquid or gas flow rates promote gas hold-up. Variations of gas hold-up as a function of liquid energy dissipation rate per unit volume of the bioreactor (E/V)L are given in Figs. 8–10. Velan and Ramanujam [3] presented a correlation for gas hold-up inside down flow jet loop reactors based on dimensionless parameters. Jamshidi [1] modified this correlation to include the effect of draft tube height and the position of the nozzle inside the draft tube as follows: ε = 1.1072 × 10−4 (Rej )1.095 (Fr)0.885
1.4Ld 1 + 2Hn
0.9
Dd Dr
0.025
(4)
Fig. 6. Influence of liquid flow rate on gas hold-up.
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Fig. 9. Influence of energy dissipation rate (E/V)L on gas hold-up. Fig. 7. Influence of liquid flow rate on gas hold-up.
Fig. 8. Influence of energy dissipation rate (E/V)L on gas hold-up.
Fig. 10. Influence of energy dissipation rate (E/V)L on gas hold-up.
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Fig. 11. Experimental and predicted gas hold-up in air–water system.
In Fig. 11, a comparison has been made between the gas hold-up experimental data and those determined using Eq. (4). The relative error is ±20%. In the above equation four dimensionless terms have been employed. The first two terms reflect the operating conditions inside the bioreactor and have a very important significance. Rej reflects the liquid flow situation in the nozzle exit port. Given the power to which Rej is raised in the above equation it can be seen that hold-up varies almost linearly with the liquid linear velocity UL , therefore, doubling of the liquid flow rate will results in a more than double increase in the gas hold-up. The second dimensionless number in the above correlation is the Froud number which has a direct relationship with gas superficial velocity. Therefore, according to the above empirical equation gas hold-up is related to the gas flow rate raised to the power 0.885. In the other words a doubling of the gas flow rate will result in a 85% increase in the gas hold-up. This means that the effect of the increase in liquid flow rate on the gas hold-up is more important than the corresponding effect of the increase of the gas flow rate. 3.1.2. Volumetric mass transfer coefficient To determine the mass transfer coefficient between the two phases, the air–water system was used. Measurement of dissolved oxygen in the fluid was performed using an oxygen meter (OXI 325-A/Set Best No. 200216 wtw) connected to an electrode. A batch of deionized water was deaerated by
bubbling argon gas. The experiment was started by passing deaerated water and air through the nozzle. Dissolved oxygen concentration was measured at intervals of 5 s. In each experiment, about 50 readings were recorded. Experiments were performed at various gas flow rates and at different locations of nozzle in the draft tube and a constant temperature of 25 ± 0.1◦ C. Oxygen saturation in water was achieved within 40–140 s, depending upon the hydrodynamic conditions. Volumetric mass transfer coefficients were calculated from the following equation: dCL = KL a(C ∗ − CL ) dt
(5)
where CL is the dissolved oxygen concentration at any time, t, C∗ is the saturation concentration of oxygen in liquid and KL a is the volumetric mass transfer coefficient. Integration of the above equation with the following boundary conditions: t = 0,
CL = Co
t = t,
CL = CL
yields the following relation: ln
C ∗ − Co = KL a t C ∗ − CL
(6)
Eq. (6) is applicable only if the flow of fluid within the bioreactor is fully back mixed. In the transient technique, the
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dynamic conditions of the oxygen electrode affect the KL a measurement. Normally, a delay or time lag is encountered in electrode response. However, the following first-order model was applied to compensate for the delay [8–10]. −t/B C ∗ − CL e−t/s Bs e − (7) = C ∗ − Co s B B −s where s and B are electrode time lag and that of 1/KL a, respectively. If t s, then Eq. (5) may be simplified as C ∗ − CL B e−t/B = C ∗ − Co B −s
(8)
Thus, the slope and intercept of the plot of ln(C ∗ − CL )/(C ∗ − Co ) versus time, gives −1/K L a = −B and ln(B/B − s), respectively; from the latter the electrode delay could be calculated. In Fig. 12, the effect of gas flow rate on the volumetric mass transfer coefficient for various locations of nozzle is shown. It is clear that increase in gas flow rate increases KL a. In addition, the position of nozzle inside the draft tube has a marked effect on KL a (thus, increasing the height of the nozzle position from the bottom of the draft tube, increases KL a). Velan and Ramanujam [4] presented a correlation for KL a inside a down flow jet loop reactors based on dimensionless parameters. Jamshidi [1] modified this correlation to include the effect of draft tube height and the position of the sprayer
Fig. 12. Influence of gas flow rate on volumetric mass transfer coefficient.
Fig. 13. Experimental and predicted mass transfer coefficient.
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inside the draft tube and came up with following correlation: Sh = 1.76(Sc)0.5 (Rej )0.99 (Fr)0.53 (Mo)−0.06 0.15 + 1.05Ld 0.71 × Hn + L d
(9)
In Fig. 13, a comparison has been made between the experimental KL a data and those estimated using Eq. (9). The degree of agreement was within ±11%. The above equation contains five dimensionless terms. The dimensionless number Rej , Fr and Mo reflects the hydrodynamic condition inside the bioreactor whilst Sc reflects the mass transfer condition inside the bioreactor. The last dimensionless term in the above equation shows the effect of the position of the nozzle in the draft tube. The exponent on Rej and Fr indicate the importance of these parameters with regards to the overall volumetric oxygen transfer coefficient. Comparison of the (4) and (9) shows
the effect of the variation of Rej and the linear liquid velocity on the overall volumetric oxygen transfer coefficient is similar to the effect of variation of these dimensionless parameters on gas hold-up whilst the effect of the variation of Fr on the volumetric oxygen transfer coefficient is appreciably less than the effect of this parameter on gas hold-up. In other words, the effect of the increase in the liquid flow rate on gas hold-up is noticeably higher than the effect gas flow rate on this parameter. This shows that in the calculation of both gas hold-up and KL a, hydrodynamic parameters play a significant role. 3.1.3. Residence time distribution (RTD) of liquid in the bioreactor To determine the RTD of liquid within the bioreactor, tracer analyses were performed. It was evident that the liquid was circulated in a closed loop through the system, while the flow of gas was continuous. It was anticipated, therefore, that a pulse of a tracer injected in the liquid flow at the entrance of the bioreactor will be evenly distributed throughout the
Fig. 14. Response at the exit of a down flow jet loop reactor subsequent to an impulse input.
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Fig. 15. Response at the exit of a down flow jet loop reactor subsequent to an impulse input.
system after a short period of time. The concentration of tracer at the bioreactor outlet might be, therefore, regarded as a measure of the hydrodynamic conditions and the pattern of flow within the bioreactor [11]. In each experiment, 1 cm3 of a known solution of potassium permanganate was injected rapidly via the liquid flow inlet to the bioreactor. Samples from the liquid phase at the outlet of the bioreactor were collected, using a circular vessel divided into 23 segments, rotating at desired speeds. By this method, successive samples were obtained at equal time intervals. The accurate determination of angular speed of the vessel was achieved using the digital Photo/Contact Tachometer DT 2236. The optical densities of samples were measured by a Pye Unicam SP 1700 UV Spectrophotometer. The relative concentrations (C/Cmax ) of samples were then calculated and RTD curves were plotted.
The tracer analysis were conducted at various liquid flow rate (5000–9000 cm3 /min) and gas flow rates (3000–50000 cm3 /min). In addition the distance between the nozzle and the top of the draft tube was varied in the range of 0–25 cm. In Figs. 14 and 15, the relative concentration (C/Cmax ) of tracer at the bioreactor outlet have been plotted versus time. Tracer analysis under different experimental conditions (various gas and liquid flow rates, position of nozzle) was performed. However, the overall patterns of RTD curves obtained remained almost unaffected. Mixing time within the bioreactor was calculated using the RTD results. In Fig. 16, the effects of inlet gas velocity and position of nozzle on mixing time are observed. It is evident that the increase in gas velocity promotes the mixing process, and hence, decreases the mixing time. Locating the nozzle at higher levels of the draft tube would enhance the
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Fig. 16. Effect of superficial gas velocity on mixing time.
circulation within the bioreactor and consequently decrease the mixing time.
4. Conclusion The following conclusions may be drawn from the present results: 1. The gas hold-up increases with increase in gas and liquid flow rates and the height of nozzle position from the bottom of the draft tube. However, the effect of gas flow rate is more pronounced. 2. The volumetric mass transfer coefficient increases with increases in gas and liquid flow rates and the height of the nozzle position from the bottom of the draft tube. 3. It may be observed from the RTD curves that the tracer concentration within the bioreactor, attains a uniform concentration after 2–6 s. Absence of long tails in the RTD curves indicates that dead regions are not present within the bioreactor. 4. The mixing time is a function of gas and liquid flow rates and the nozzle position in the draft tube. Increasing the
fluid flow rates and the height of nozzle location in the draft tube, reduces the mixing time. References [1] A.M. Jamshidi, Chemical engineering, Ph.D. Thesis, Amirkabir University of Technology, Tehran, Iran (2000). [2] C.A.M.C. Dirix, K. Van der wiele, Chem. Eng. Sci. 45 (8) (1990) 2333–2340. [3] M. Velan, T.K. Ramanujam, Chem. Eng. Sci. 47 (9–11) (1992) 2871–2876. [4] M. Velan, T.K. Ramanujam, Can. J. Chem. Eng. 69 (1991) 1257–1261. [5] N. Rabiger, A. Vogelpohl, Chem. Ing. Tech. 55 (6) (1983) 486–487. [6] Karl Schugerl, Bioreaction Engineering, Vol. 2, Wiley, New York, 1991. [7] K. Kulkarni, Y. Shah, A. Schumpe, Chem. Eng. Commun. 24 (1983) 307–337. [8] M. Nakanoh, F. Yoshida, Ind. Eng. Chem. Process Des. Dev. 19 (1980) 190. [9] S.A. El-Temtamy, S.A. Khalil, A.A. Nour-El-Din, A. Gaber, Appl. Microbiol. Biotechnol. 19 (1984) 376. [10] M.Y. Chisti, M. Moo-Young, Biotech. Bioeng. (1988) 487–494. [11] M. Sohrabi, A.M. Jamshidi, J. Chem. Tech. Biotechnol. 69 (1997) 415–420.