Rotating Biological Contactors: A Review - American Chemical Society

Ind. Eng. Chem. Res. 2003, 42, 2035-2051

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REVIEWS Rotating Biological Contactors: A Review A. W. Patwardhan* Institute of Chemical Technology, University of Mumbai, Matunga, Mumbai 400 019, India

Rotating biological contactors (RBCs) offer an alternative treatment technology to the conventional activated sludge process. The principal advantages of the RBC stem from the fact that the interfacial area generated is very high and practically independent of the speed of rotation, unlike in the activated sludge process. In the present work, the process design aspects of RBCs are reviewed. The published literature on various aspects such as media employed in RBCs, hydrodynamic characteristics, power consumption, and mass-transfer characteristics is reviewed. The various models proposed for the design of RBC systems are reviewed. A stepwise procedure is given for the process design of an RBC system. The RBC system is compared with the other wastewater treatment facilities such as the activated sludge process and trickling filters in terms of its oxygen-transfer efficiency. Contents 1. 2. 3. 4. 5. 6.

Introduction Media for RBCs Hydrodynamics of RBCs Power Consumption of RBCs Mass-Transfer Characteristics Structure, Transport, and Reactions in Biofilms 7. Effects of Operating Parameters on the Performance of RBCs 7.1. Effect of Speed of Rotation 7.2. Effect of Tank Volume to Surface Area Ratio 7.3. Effect of Staging 7.4. Effect of Organic and Hydraulic Loading Rates 7.5. Effect of Temperature 7.6. Effect of Recirculation of the Effluent 8. Mathematical Models for Performance Prediction 9. Process Design Aspects 10. Comparison with Other Wastewater Treatment Systems 10.1. Surface Aerators 10.2. RBC Systems 10.3. Trickling Filters 11. Conclusions 12. Suggestions for Future Work 13. Nomenclature 14. Literature Cited

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1. Introduction Rotating biological contactors (RBCs) are widely employed in wastewater treatment. RBCs essentially * Tel.: 91-22-414 5616. Fax: 91-22-414 5614. E-mail: awp@ udct.ernet.in.

consist of a set of disks mounted on a horizontal shaft. The disks are partly submerged in the wastewater and partly exposed to the air. The entire rotor assembly consisting of a shaft with the attached disks is rotated at 1-10 rpm with the help of a motor or compressed air.1 The rotation of the disks generates a film of liquid on the surface of each disk. As the disk rotates, the film is exposed to air that enables aeration of the wastewater and allows aerobic organisms fixed on the surface of the disk to degrade the waste. The set of disks on which the biomass grows is called the media. Thus, in principle, these contactors enable contact between the gas phase and the liquid phase. Shear forces exerted on the biomass during the passage of the disk through the water causes the excess biomass to be stripped from the media. This prevents clogging of the media and maintains a constant microorganism population on the media. The media serve the following functions: (i) provide surface area for the development of a fixed biological culture, (ii) enable contact of the wastewater with the air, thereby aerating the wastewater. Typically, a set of disks is held together in a unit, and several such units can be arranged in series or parallel. The biofilm and the liquid film generated by the disk rotation is shown in Figure 1A. The principal advantages of RBCs stem from the fact that the interfacial area generated is very high and practically independent of the speed of rotation, unlike the case for the activated sludge process. The power consumption is low mainly because (i) the disks can be made thin so that they offer little resistance to movement through the wastewater and (ii) the speed of rotation is low. The low power consumption results in lower operating costs. Greaves et al.2 compared the operating costs of RBC plants with those of extended aeration systems and showed that the savings in operating costs of RBC systems over conventional extended aeration systems increase significantly with increasing size of the plant. The effects of various operating parameters on RBC performance were compiled by Antonie1 and Brenner et al.3 Using pilot-plant data from various installations,

10.1021/ie0200104 CCC: $25.00 © 2003 American Chemical Society Published on Web 04/11/2003

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Ind. Eng. Chem. Res., Vol. 42, No. 10, 2003

Figure 1. (A) Overall setup of an RBC unit. The biofilm and liquid film generated on the media. (B) Typical media used in RBCs.

they proposed various criteria for the process design of RBC systems. Their reports suggest that the important design criteria include the hydraulic loading (flow rate per unit time per unit surface area covered by the biofilm), series/parallel arrangement of the RBC units, and rotational speed (tip velocity) of the media. The hydraulic loading (flow rate per unit surface area in the reactor, HL) can be related to the residence time (ratio of the volume of the reactor to the flow rate, RT) within the reactor through the void volume in the media and the interfacial area provided. If a is the interfacial area provided by the media and is the void volume, then the residence time and the hydraulic loading can be related by

(RT)(HL) ) /a

(1)

Antonie1 recommended that the hydraulic loading should be kept between 1 and 10 gpd/ft2 depending on the inlet BOD and the desired degree of removal. Commonly, RBC systems consist of several units placed in series. This allows for the development of specific microbial cultures in the different stages of the media. It also makes the overall flow pattern approach plug flow. Antonie1 and Brenner et al.3 recommended that the RBC unit should be made into a four-stage operation and that the peripheral velocity should be kept about 0.3 m/s. Such criteria were determined through experimentation on a pilot-plant scale. However, scale-up based on such criteria should be done with caution. For example, if it is desired to scale-up from say 1 to 100 m3/day wastewater flow, then keeping the hydraulic loading would require that the number of disks increase by a factor of 100. In addition, the criterion of equal tip speed would suggest that, if the disk diameter were increased by a factor of 5 upon scale-up, then the rotational speed would have to be reduced by a factor of 5. The power requirements on the larger scale would then increase by a factor of 2500, and the power per unit volume would increase by a factor of 25. The liquid hold-up on the media would be about 2500 times greater (eq 4), and the fractional liquid volume on the disk surface would increase by a factor of 25. The average contact time would increase by a factor of 5 (eq 12).

Alternatively, if scale-up were done at equal hydraulic loading and equal number of disks, then the disk diameter would have to be increased by a factor of 10. In addition, if the tip speed were maintained the same, then the rotational speed would need to be reduced by a factor of 10. The power requirements would go up by a factor of 100, and the power per unit volume would remain the same. The liquid hold-up on the media would increase by a factor of 100, and the fractional liquid volume on the disk surface would remain the same. The average contact time would increase by a factor of 10. Clearly, scale-up based on such criteria would dramatically influence the hydrodynamic characteristics within the reactor. It is likely that a different level of power per unit volume on a larger scale would alter the stresses experienced by the microorganisms, as well as the biomass loading and the thicknesses of the biofilm and the liquid film. As a result, the performance is likely to be affected. The drawbacks of such global design criteria (fixed tip speed, hydraulic loading between 1 and 10 gpd/ft2) are that the important aspects such as hydrodynamics, mass transfer, kinetics of biodegradation, etc., are not considered at all. The generalized design criteria also might not result in an economically optimum design. Therefore, it was thought desirable to critically review the literature on rotating biological contactors, especially the hydrodynamics and masstransfer characteristics of these systems. It was also thought necessary to compare the performance of RBC systems with that of other wastewater treatment facilities such as those using the activated sludge process employing surface aerators, trickling filters, etc. Mba et al.4 reported that, even though RBC units offer several advantages, they are susceptible to mechanical failure due to poor design, low-frequency corrosion fatigue, and microbial corrosion. This conclusion was drawn by analyzing the causes of failure of more than 250 RBC units. A mechanical failure of the RBC shaft implies a high replacement cost because the failure of the shaft also causes damage to the media. The bearing can fail due to lack of lubrication (which can be due to lack of grease, grease contamination, incorrect type of grease, etc.). The supporting structure of the RBC media is likely to fail due to fatigue. Once a portion of this support structure fails, additional load is created on the remaining structure, which can result in a sudden failure of the entire set of media. They concluded that, in many RBC units, failure is due to fatigue, which implies that this factor was not properly taken into account during mechanical design. Microbial corrosion was attributed to the excess growth of microorganisms and the growth of incorrect types of organism. This could be controlled by reducing the substrate loading or by increasing the speed of rotation. Microbial corrosion occurs as a result of the production of corrosive metabolites such as sulfides, organic acids, etc. Mba et al. suggested several guidelines for mechanical design so as to reduce mechanical failures. 2. Media for RBCs Over the years, various types of disk media have been developed. Their primary functions are to provide surface area for biological growth, bring about contact between the wastewater and the microorganisms, and achieve oxygen transfer from air into the film of wastewater and biofilm. About a century ago, in the 1900s, RBC media consisted of a cylinder with wooden slats. However, the

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Figure 2. Schematic representation of packed-cage RBC developed by Wanner et al.6

wooden slats were prone to clogging, so they were replaced by metal disks in the 1930s. In the 1950s, the metal disks were replaced by expanded polystyrene disks. These disks were about 1 cm thick and were placed 3 cm apart. This arrangement could provide a surface area of only about 25 m2/m3 of reactor. At 50% submergence, the surface area would only be 50 m2/m3 of liquid. In the 1970s polyethylene disks were introduced so as to reduce the fabrication costs. An added advantage of polyethylene is that it can be easily fabricated into corrugated sheets having thicknesses between 1 and 2 mm. Various corrugation patterns have been reported in the literature. These depend on the manufacturer. Some typical configurations are shown in Figure 1B. These corrugations increase the surface area for contact between the wastewater and the air. Because the media are corrugated and the disks are closely packed, the identity of the disks is lost. Therefore, the various types of media are classified on the basis of surface area provided. The media are typically classified into two types: standard and high-density. The standard type provides 100 000 ft2 area on a 27-ftlong shaft with a diameter of 12 ft, whereas the highdensity type provides about 150 000 ft2 area on a 27 ft long shaft with a media diameter of 12 ft. If these value are converted to the usual notation, this means that the surface area is about 125 m2/m3 of reactor or 250 m2/ m3 of liquid for high-density media. This compares well with the typical values of surface area in conventional reactors5 such as packed columns (25-200 m2/m3 of reactor), bubble columns (25-500 m2/m3 of reactor), and stirred reactors (100-1000 m2/m3 of reactor). Wanner et al.6 developed an RBC system consisting of a cage filled with random packings, which is shown schematically in Figure 2. The cage was constructed with wire mesh netting and was filled with cylindrical plastic elements. Wanner et al.6 reported that their entire assembly provided a contact area of 10.5 m2. For the dimensions given in their paper, the interfacial area is 100 m2/m3 of reactor volume, which is the same as that of standard-density media (Figure 1B). Ware et al.7 reported that randomly packed high-voidage plastic media can be used as a substitute for the conventional high-density media. They carried out experiments with conventional packing media such as rings on a large scale, including 2.2- and 3-m-diameter packed cages. They observed that a completely packed cage unit with little or no movement of the packing can lead to excessive growth, which can ultimately lead to short circuiting and loss of treatment efficiency. They further reported that a packed cage with large allowable movement for the packing can lead to excessive shear due to rubbing/tumbling of packings over each other and that this can destroy the biofilm completely.

Figure 3. Schematic representation of RBC developed by McManus.8

McManus8 reported the use of polyethylene/polypropylene random loose media. The loose nature of the media allows the media to tumble during the rotation of the system, as shown schematically in Figure 3. It has been claimed that such systems aerate wastewater more efficiently and enable self-cleaning. Nahid et al.9 carried out experiments with an RBC system containing polypropylene pall rings as packing material. To date, there has been no clear comparison between the various types of media that would enable the process design engineer to select a particular type of media. From the above discussion, it is clear that RBC systems have evolved considerably from the original design of several rotating disks into a unit in which the disks tend to lose their identity and behave as a single unit consisting of random/structured packings. The whole RBC system can then be visualized as a horizontally placed rotating packed column, but with the following major differences: (i) Instead of a vertical packed column, the packed section is kept horizontal. (ii) The liquid-phase distribution is ensured by rotation of the packed section, rather than by the use of distributors as in conventional packed columns. This also means that the wetting of the entire packed section can be achieved without considerations such as the minimum wetting rate in conventional packed columns. (iii) The gas phase need not be supplied by a compressor, as the rotation of the packed section ensures good contact between the gas and the liquid phases. The chances of channeling of the gas and liquid phases are minimized. (iv) Because the gas phase need not be supplied with the help of a compressor, the issues of gas-phase pressure drop, flooding, etc., are eliminated. (v) Thus, the RBC system need not be restricted to wastewater treatment but can be used for any gas-liquid contacting operation. In fact, rotating disk contactors (RDCs) are already being used for blood oxygenation and as evaporators.5 If the headspace can be closed completely, the entire system can be operated as a dead-end system. Dead-end systems are particularly advantageous for

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carrying out gas-liquid or gas-liquid-solid reactions where internal recycling of the unreacted gas is desirable for the reasons of cost, safety, or corrosion.10 (vi) The only major difference is that the liquid hold-up in RBC unit is much larger than that in conventional packed columns; therefore, it might not be suitable for the processing of heat-sensitive materials. (vii) The extent of axial mixing in RBCs might be larger than that in conventional packed columns. In fact, recently, Gai et al.11 employed a rotating trickle-bed reactor for biological gas cleaning. This system consisted of a structured packed bed rotating on a vertically placed shaft, with the liquid and gas being fed from the top. Recommendations. It is necessary to compare the various types of media reported in the literature in terms of the interfacial area offered, mass-transfer coefficients, and power consumption. This will enable the process design engineer to choose the most appropriate type of media. It would be worthwhile to employ structured packings in rotating biological contactors and to compare the efficiency of such media to that of other media.

Figure 4. Schematic representation of film and boundary layer formation on the disk surface.

3. Hydrodynamics of RBCs Most RBC hydrodynamic studies have pertained to the use of plain disks in RBC systems. When the disk rotates, it carries with it a film of liquid in the gas phase. Oxygen transfer occurs in this film by diffusion. If a biofilm is also present, oxygen diffuses further into the biofilm, where it is used for the degradation of the waste from the wastewater. The liquid film formed on the disk surface mixes with the bulk liquid, thereby aerating the bulk liquid. Thus, the important hydrodynamic parameters are generation of the film, film thickness, flow velocities within the film, mixing of the film with the bulk liquid, etc. Yamane and Yoshida12 and Bintanja et al.13 assumed that the liquid film entrained on the disk surface is stripped off and completely mixed with the bulk liquid as soon as the disk re-enters the liquid. Zeevalkink et al.14 correlated the average film thickness developed on the disk using the equation

δ)

4 2ν 1/2 1/2 vC 15 g

( )

(2)

Suga and Boongorsrang15 assumed that, when the disk re-enters the liquid, a boundary layer is developed on the disk surface, through which oxygen transfer to the bulk liquid occurs. The film and the boundary layer formed on the disk surface are shown schematically in Figure 4. They also carried out visualization studies of the process of mixing of the film with the bulk liquid. It was observed that the liquid film was partly stripped off from the disk and mixed with the bulk liquid when the disk re-emerged from the liquid rather than when it entered. This process is depicted in Figure 5. The photographic evidence presented by them is clearly contrary to the assumptions made by earlier workers, namely, Yamane and Yoshida12 and Bintanja et al.13 Vaidya and Pangarkar16 accounted for the local variation in the film thickness due to the action of forces such as centrifugal forces, gravity, surface tension, etc. The local film thickness was estimated from the correlation

Figure 5. Stripping of the liquid film and its subsequent mixing within the liquid.

developed earlier17

TO (1 - TO2)2/3

) 0.944Ca1/6

(3)

where TO is the dimensionless film thickness and Ca is the capillary number. If the denominator is close to unity, then the above equation reduces to the film thickness being proportional to (velocity)0.66, whereas the exponent is 0.5 from the correlation of Zeevalkink et al.14 They reported that, for water, at disk tip velocities in the range of 0.2-3 m/s, the maximum film thickness varies between 60 and 150 µm. Knowing the film thickness, they also calculated the volumetric holdup of liquid attached to the disk surface using the correlation

V ) 8.4 × 10-4ω0.63R2.63

(4)

Within the film, they solved the simplified Navier stokes equations to yield the radial and tangential velocities of the liquid phase. The axial velocity component (perpendicular to the disk) was assumed to be 0. The Navier-Stokes equation in the radial direction was simplified by considering that only the centrifugal, gravitational, and shear (τrz) forces are important and other terms can be neglected. In the tangential direction, it was assumed that the gravity, shear (τzθ), and convection were important. The numerical solution of the simplified equations showed that vθ is practically


independent of r and z and that the values of vr are small as compared to vθ. This model was extended for any level of submergence and for non-Newtonian liquids by Vaidya and Pangarkar.18,19 Sassi et al.20 investigated the hydrodynamics of the RBC system using the characteristic time scales for various processes. They defined various time scales such as mean residence time, time constant for mixing, time constant for substrate degradation, time constant for oxygen uptake, time constant for oxygen transfer, etc. The values of these time constants were determined from experimental investigations on a system consisting of 25 plexiglass disks having a diameter of 0.2 m. From their experimental investigations, it was observed that, at low disk rotational speeds and low liquid flow rates, the time constant for mixing was much larger than the mean residence time. This would mean that the overall system would behave close to plug flow. In contrast, at high disk rotational speeds, the time constant for mixing became comparable to the overall residence time, which would mean that the overall system would behave between a system with plug flow and a well-mixed tank. Banerjee21 investigated the overall flow characteristics of an RBC system. He carried out extensive experimental measurements and investigated the effects of the liquid flow rate, rotational speed, and degree of submergence. His experimental setup consisted of 60 flat polystyrene disks with a 0.41-m diameter. The speed of rotation was varied from 1 to 15 rpm, the liquid flow rate was varied in the range of 0.04-22 m3/day, and the submergence was varied between 30 and 60%. The extent of backmixing in the RBC unit was calculated by measuring the outlet concentration of the tracer as a function of time. The overall RTD curve obtained in this manner was fitted with a plug flow with axial dispersion model. The dispersion number was found to vary over the range 0-8 for the entire range of their experimental investigations. The following correlation was proposed to estimate the dispersion coefficient

( ) ( )

DL d hω 2 ) 0.32 UL g

0.57

d hU ν

0.75

(5)

Recommendations. It can be concluded that all of the hydrodynamic studies performed to date pertain to the use of flat disks in the RBC system. The hydrodynamic characteristics such as the liquid film thickness and the extent of axial mixing, for the various types of media mentioned above have not been investigated. As a first approximation, one could assume that each unit of an overall RBC system behaves as a CSTR. However, detailed experimental investigations on the media described above need to be carried out to quantify the axial mixing. 4. Power Consumption of RBCs Antonie et al.22 presented the power consumption data for 150 disks, each of 3-m diameter. From their data, the power number for each disk can be calculated using eq 6 below. The power number computed in this manner comes to about 0.2 per disk. Poon et al.23 indicated that a 3.2-m-diameter RBC system providing 24 820 m2 of surface area requires about 11 kW of power. However, from these data, the power number cannot be computed, as the speed of rotation was not given.

Fujie et al.24 correlated the power drawn with the help of a power number by analogy with impeller power number in stirred reactors

NP )

P FLN3D5

(6)

Under laminar flow conditions, the power number was assumed to be inversely proportional to the disk Reynolds number based on the disk diameter and the distance between the disks

NP )

R , where Re ) NDdFL/µ Re

(7)

Using the above relations, the power requirement per unit surface area can be calculated as

P/A ) λ1N2D2, where λ1 )

4Rµ πd

(8)

The above expression is valid in the laminar regime only, up to a peripheral disk velocity of 0.5 m/s. Under turbulent flow conditions, the disk power number can be considered to be independent of the Reynolds number. Under these conditions, the power requirements per unit surface area can be calculated as

P/A ) λ2N3D3, where λ2 )

4NPFL π

(9)

Fujie et al.24 analyzed the data published by the U.S. EPA and showed that λ1 ) 0.86 × 10-5 kW min2 m-4 and λ2 ) 0.052 × 10-5 kW min3 m-5. If these data are reanalyzed in terms of the power number, it turns out to be about 0.1. The power consumption data for various RBC installations equipped with mechanically driven units have been compiled by Brenner et al.3 and Gilbert et al.25 Their data compilation is exhaustive, consisting of results from about 20 locations (105 RBC units). Their data show that the RBC power requirement ranges from 1.21 to 3.12 kW/shaft. Their measurements included standard and high-density media, with average shaft lengths of between 20 and 25 ft and media diameters of 11-12 ft. Taking an average value of the extensive set of available measurements shows that, at a speed of 1.5 rpm and an average surface area of 125 000 ft2, an RBC draws an average power of 2.30 kW. When power numbers are calculated for their data, they turn out to be in the range 0.05-0.15. Recommendations. It would be highly desirable to compare the power consumption characteristics of the various types of packings mentioned above. This will help the process design engineer in selecting the appropriate type of media. In the absence of such data, it seems reasonable to assume a power number of about 0.1 per disk. 5. Mass-Transfer Characteristics Yamane and Yoshida12 were among the first researchers to theoretically analyze the mass-transfer aspects. They modeled the oxygen transfer to the liquid film

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Figure 6. Schematic representation of the contact between gas and liquid phase in an RBC.

(thickness ) δ) on the disk surface as

∂C ∂2C ) Dm 2 ∂t ∂x

(10)

with the boundary conditions

t ) 0, 0 < x < δ, C ) Ci t > 0, x ) δ, C ) C*

(11)

∂C )0 ∂x

t > 0, x ) 0,

Solution of eq 10 gave the concentration profile of oxygen within the film. Using the calculated concentration profiles, the flux of oxygen was calculated. Using this, the values of mass-transfer coefficient were calculated. The mass-transfer coefficients predicted in this manner were compared with those obtained from the considerations of the average contact time and Higbie’s penetration theory. It was observed that these two approaches gave practically the same result for average contact times of less than 0.2 s but deviated substantially at higher contact times. The discrepancy could be attributed to the fact that different fluid elements have different contact times. The contact of a particular fluid element with air is shown schematically in Figure 6. The contact time can be calculated from the peripheral distance and the angular velocity at that radial location

tC )

R-H π-φ 1 1 ) 1 - cos-1 Nπ N π r

(

)

(12)

From the above expression, it can be clearly seen that the contact time is a function of the radial distance. However, if the disk is filled to 50% submergence, then H becomes equal to R, and the angle φ becomes π/2. The local contact times were integrated to give the average contact time. The average contact time as a function of the H/R ratio is shown in Figure 7. Using this average contact time, the mass-transfer coefficient can be estimated with Higbie’s penetration theory using the equation

kL ) 2

x

Dm πth

Figure 7. Relation between the average contact time and the submergence ratio, H/R, estimated by Yamane and Yoshida.12

Ravetkar and Kale26 assumed a uniform film thickness on the disk surface. However, they considered only the average contact time, which was calculated from the tip velocity and the circumference of the disk. In addition, they considered that the disk surface wetted by the liquid depends on the submergence of the disk. The interfacial area exposed to the gas can then be written as (Figure 6)

[

Ad ) 2πR2

(

(14)

In addition to the oxygen transfer from the gas to the liquid film on the disk surface, they considered that oxygen transfer occurs (i) at the disk periphery and (ii) on the surface of the liquid between the disks. The total interfacial area was written as a sum of these three areas. Using this expression, the interfacial area per unit contactor volume was calculated. This result was then used to calculate the overall mass-transfer coefficient (kLa). This expression for kLa was differentiated with respect to the submergence, and it was observed that the maximum value of the mass-transfer coefficient occurred at 42% submergence of the disk. Their experimental setup consisted of eight acrylic disks having 125and 175-mm diameters. It was observed that the interfacial area per unit volume was independent of the speed of rotation of the disk and that it increased with a reduction in the submergence from 50 to 40%. All of the data on the mass-transfer coefficient were correlated using the equation

kLD ) 1.59(D2N/Dm)0.5 Dm

(15)

Kim and Molof27 conducted experimental measurements of the overall mass-transfer coefficient. Their experimental setup consisted of 1-4 flat plastic disks having diameters in the range of 0.15-0.6 m. The disk spacing was varied in the range of 6-70 mm. The disk rotation speed was varied from 5 to 85 rpm. The oxygen concentration was measured with a dissolved oxygen meter. The volumetric mass-transfer coefficient was correlated using the equation

kLa ) 0.0011 (13)

)]

(π - φ) sin 2φ 1 + cos 2φ + π 2π 2

(

)

ω1.5D0.5 d

0.732

(16)

The term in parentheses was defined as the volume


renewal number. This is not a dimensionless number. The units of ω should be rotations per minute, whereas those of D and d should be inches; thus, the calculated kLa has units of inverse minutes. Suga and Boongorsrang15 assumed that boundary layer formation takes place on the disk surface during its travel through the liquid. The oxygen transfer occurs through this boundary layer. This process is depicted in Figure 4. The oxygen transfer occurring through the liquid film exposed to air was modeled using the diffusion equation in a manner similar to Yamane and Yoshida.12 The oxygen transfer through the boundary layer was assumed to be purely by diffusion. The flux of oxygen in the liquid film exposed to air was expressed as

R2 dθ 2

dNC ) k′L(C* - C′)

(17)

The volumetric flow rate of the liquid due to its entrainment on the film was given as

Q ) NπR2δ

(18)

Writing the overall oxygen balance for the film exposed to air gave

ln

C* - C′1 K′L ) C* - C′2 2Nδ

(19)

A similar balance for the oxygen transfer through the boundary layer in the liquid phase gave

-ln

KL C′′1 - C ) C′′2 - C 2NδBL

(20)

These equations were solved to obtain the values of the oxygen concentration in the liquid film just after and just before disk-water contact. Using these concentrations, an overall volumetric mass-transfer coefficient was calculated for the disk. The oxygen concentrations predicted from the above model agreed well with the experimental measurements made using 10 PVC disks having a diameter of 15 cm. Vaidya and Pangarkar16 reported that the values of vθ obtained by solution of the Navier-Stokes equation within the liquid film are practically independent of r and z and that the values of vr are small as compared to vθ. In view of this, the conservation equation for the oxygen can be written as

vθ ∂C ∂2 C ) Dm 2 r ∂θ ∂z

(21)

The above equation was solved numerically for the exposed portion of the disk and the concentration profiles were calculated. Using the concentration profiles, the volumetric rate of oxygen absorption was calculated and correlated as

RC ) 2.15 × 10-8

( ) 2πDm ωZ

2

-0.259

R1.71

(22)

The parameter Z in the above equation represents the maximum film thickness at the tip of the disk. For water, the value of Z was correlated as Z ) 7.05 × 10-4(Rω)0.63.

This basic model was later extended18,19 to take into account the effect of submergence of the disk and the non-Newtonian nature of the liquid phase. This analysis was done only for a single disk and for physical transfer of oxygen. To extend this model to an entire RBC unit, the following issues need to be addressed: (i) the presence of other disks, (ii) transfer of oxygen along with its simultaneous consumption by microorganisms, and (iii) the kinetics of degradation of the waste in the wastewater. Paolini28 investigated the effect of biomass on the oxygen-transfer characteristics of RBC systems. Their experimental setup consisted of 11 flat polystyrene disks having a diameter of 0.28 m. The rotational speed of the disk was varied in the range of 3-25 rpm. It was observed that, in the presence of microorganisms, there is a very large enhancement in the rate of mass transfer. At low speeds (less than 5 rpm), the enhancement was almost 10 times; however, as the speed increased, the enhancement factor was found to decrease. Nishidome et al.29 developed a microelectrode for the measurement of dissolved oxygen (DO) profiles in the liquid and biofilm developed on the disk surface. Measurements of the dissolved oxygen concentration enabled the computation of the oxygen flux in the liquid and the biofilm. Boumansour and Vasel30 carried out detailed investigations on the mass-transfer rate and its enhancement due to the presence of the biofilm. Their experimental setup consisted of nine disks having diameters of 0.25 m. The speed of rotation was varied in the range of 6-35 rpm. The mass-transfer coefficient was calculated by measuring the dissolved oxygen concentration in the liquid phase. They showed that their experimental data as well as the data of previous workers could be correlated using the following expression

( ) ( ) ()

kLD ND2F ) 2.673 Dm µ

0.769

N2D g

0.135

H R

0.865

(23)

Recommendations. Thus, it can be concluded that all of the reported mass-transfer studies pertain to the use of flat disks in RBC systems. Experimental investigations of the mass-transfer characteristics of various types of media described above would help in quantifying the mass-transfer rates for those kinds of media. The detailed models proposed by Vaidya and Pangarkar16,18,19 take into account the flow pattern and the variation of the liquid film thickness on the disk surface. Comparing the correlations proposed by various workers, it can be seen that the correlation of Ravetkar and Kale26 predicts kL ∝ N0.5D0, the correlation of Vaidya and Pangarkar16 predicts kL ∝ N0.585D0.326, and the correlation of Boumansour and Vasel30 predicts, kL ∝ N1.039D0.673 at constant H/R ratio. Thus, there seems to be a great deal of difference in the exponents predicted by these three workers. The estimation of Vaidya and Pangarkar16 is based purely on mathematical modeling and is not supported by experimental evidence, whereas the other two correlations are based on experimental measurements. Boumansour and Vasel30 showed that their correlation is also valid for data from other previous workers, and hence, their correlation would be expected to predict mass-transfer coefficients more accurately. No further work has been published on the mass-transfer characteristics of packed-cage RBC systems. In view of this fact, experiments were carried out

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with the following objectives: (i) experimentally measure the mass-transfer coefficients over a wide range of speeds of rotation and depths of submergence of the disk and check the applicability of the correlation of Boumansour and Vasel30 and (ii) compare the masstransfer characteristics of disks and packed-cage rotating biological contactors. The experimental measurements were carried out in an RBC system consisting of a semicircular trough (with a straight portion at the top) having a 0.3-m diameter and a 0.5-m length. The disks were 0.25 m in diameter and had a thickness of 2 mm. The system consisted of 21 disks. The disks as well as tank were made of transparent acrylic to enable visual observations. The speed of rotation of the disks could be varied over a wide range by means of belt and pulley system. The drive motor was 1/4 hp. All of the experiments were carried out with air-water only, and no biofilm was allowed to form on the disks. This was done to allow measurement of the mass-transfer coefficient in the absence of reaction of the dissolved oxygen with the substrate. Thus, the interference of chemical reaction and subsequent enhancement in mass transfer was prevented. A typical experiment consisted of taking a certain volume of tap water (depending upon the level of submergence desired) in a separate tank. The dissolved oxygen was measured in this tank with a dissolved oxygen electrode. Depending on the dissolved oxygen concentration and the volume, the stoichiometric requirement of sodium sulfite (to react with all of the dissolved oxygen) was calculated. A slight excess (typically 10% excess) of sodium sulfite and a small quantity of cobalt sulfate were added to the tank and mixed well. The sodium sulfite reacted with the dissolved oxygen in this tank in the presence of the Co2+ ions, and as a result, the dissolved oxygen concentration decreased with time. This decrease was monitored with a dissolved oxygen probe. When the dissolved oxygen reached 0, the solution was carefully transferred to the RBC system. The dissolved oxygen probe was connected to the bottom of the RBC system, and disk rotation was started at the desired speed. Because of the rotation of the disks, oxygen was transferred from the gas phase to the liquid film and was mixed with the bulk liquid. As a result, the dissolved oxygen in the system started increasing. The change in dissolved oxygen in the system was monitored with time, and the mass balance equation for dissolved oxygen was written as

dC ) kLa(C* - C) dt

(24)

The above equation was integrated, which enabled calculation of kLa as

kLa )

(

)

C* - Ci 1 ln t C* - C

(25)

In this manner, the mass-transfer coefficient was calculated at various speeds of rotation and for various levels of submergence. To compare the mass-transfer characteristics of a cage-type rotating biological contactor, a wire-mesh cage was constructed having the same diameter and (a length of 0.4 m) and was filled with 1-in. plastic pall rings. The results of these measurements are shown in Figure 8. Figure 8 shows that, for the RBC system consisting of disks, an increase in the speed of rotation (at a given

Figure 8. Mass-transfer coefficients of disk and packed-cage systems: O, 53 rpm disk system; 4, 40 rpm disk system; 0, 35 rpm disk system; ], 23 rpm disk system; 9, 28 rpm packed-cage system.

Figure 9. Mass-transfer coefficient data obtained in this work expressed in terms of the correlation of Boumansour and Vasel.27

level of submergence) leads to an increase in the value of the mass-transfer coefficient. Further, at a given speed of rotation, a reduction in the submergence leads to an increase in the mass-transfer coefficient. A reduction in the submergence of the disk results in more area being made available for film formation; as a result, the mass-transfer coefficient increases. However, when the submergence was reduced to 25%, the mass-transfer coefficient was found to decrease. A comparison of the mass-transfer coefficients for disk and packed-cage systems shows that, at similar speeds of rotation, the mass-transfer coefficient for the packed-cage system is about 3-4 times higher that that of a system consisting of disks. To test the validity of the correlation of Boumansour and Vasel,30 the data from the present work was plotted with the help of their correlation. The results are shown in Figure 9. From this figure, it can be seen that the correlation of Boumansour and Vasel30 is able to predict the mass-transfer coefficients well. The only difference is that the value of the constant was found to be larger than that reported by Boumansour and Vasel.30 This could be because of the differences in the methods of measurement, locations of the measurements, natures of the disk surfaces, impurities present in tap water, etc. 6. Structure, Transport, and Reactions in Biofilms The transport of oxygen and the substrate through the liquid film (Figure 1A) makes them available for biodegradation in the biofilm. The substrate is degraded depending on the rate of metabolism of the microorganisms. This, in turn, depends on the type of organisms present, biomass loading, substrate concentration, dis-


solved oxygen concentration, etc. The microorganisms grow by degrading the substrate and, in turn, generate cell mass. If the entire surface area of the disk is covered with biofilm, then the generated cell mass can be accommodated only by a corresponding increase in the biofilm density or thickness. The hydrodynamic shear experienced by the biofilm causes detachment of excess growth. Thus, the thickness of the biofilm is determined by the ability of the microorganisms to generate cell mass (YX/S), their ability to adhere to each other and to the disk surface (this will govern the biofilm density), and the level of hydrodynamic stress that they can withstand. Truelear and Characklis31 proposed that biofilm development on any surface exposed to a fluid is the net result of several physical, chemical, and biological processes, namely, (i) transport and adsorption of substrate to the surface, (ii) transport of microbial cells to the surface, (iii) attachment of the microorganism to the surface, (iv) microbial transformations (growth and exopolymer production) at the surface resulting in the production of the biofilm, and (v) partial detachment of the biofilm caused by fluid shear stresses. The first two processes are physical in nature. The attachment of the microorganisms to the surface occurs through the production of a polysaccharide binding material. Biofilm production results from the reproduction of microbial cells and the production of extra cellular protein. In this section, only a brief review of the literature pertaining to these aspects in rotating biological contactors and rotating drum/cylinder reactors is presented. Bishop32 has provided a detailed review of the biofilm structure and its influence on the transport properties. A number of papers dealing with the diffusivity of oxygen in biofilms and the effects of various parameters are also available (e.g., deBeer and Stoodley33 and Leewandowski et al.34). Truelear and Characklis31 carried out experiments with two concentric cylinders. The inner cylinder was rotating, and the outer cylinder was stationery. Their experimental measurements included the biofilm thickness, the substrate (glucose) removal rate, the torque (measure of the shear stresses), etc. It was observed that the substrate removal rate increased with increasing biofilm thickness but only up to a certain thickness. This critical value was called the active thickness. The value of the active thickness was found to depend on the substrate concentration in the reactor. The biomass production rate was also found to depend on the glucose concentration and the glucose loading rate. On the basis of this work, they correlated the specific biomass production rate assuming Monod-type kinetics. The values of the saturation concentration and the growth rate were calculated from the experimental results. They also reported that the different microbial species were preferentially selected in the biofilm depending on the substrate concentration. At low concentrations, filamentous organisms were found to prevail. They also found that the biofilm detachment rate increased with increasing mass of biofilm but was independent of the substrate loading rate. The biofilm detachment rate was found to increase with increasing speed of rotation. Christensen et al.35 investigated the biofilm structure in a concentric cylinder reactor. They observed an exponential increase in the rate of substrate removal with increasing biofilm thickness. These experiments were not carried out over a wide range of biofilm thicknesses. It is likely that these experiments were

below the active thickness reported by Truelear and Characklis.31 It was also observed that, during this period, the biofilm density increased substantially. When visual observations were made, the biofilm was found to be essentially heterogeneous, with large variations in the biofilm density in different regions. They proposed that, when the biofilm density is low, the transport of substrate through the biofilm might become the controlling factor. At high biofilm densities, the degradation can be considered to be occurring only in a thin region at the surface of the biofilm. Arvin36 investigated the kinetics of degradation of chlorinated hydrocarbons using a rotating cylinder apparatus. He determined the specific consumption (yield coefficient) of oxygen and the biomass per unit mass of methane oxidized. The yield coefficient of oxygen was found to be between 2 and 3.6 mg of O2/mg of CH4, depending on the metabolic state of the microorganisms. The yield coefficient for biomass was found to be around 0.45 mg of biomass/mg of CH4. The rates of degradation of di- and trichloroethanes and chloroethenes were found to be first-order with respect to the substrate concentration. However, the concentration range studied in this investigation was quite small (only up to 4 mg/L). Zahid and Ganczarczyk37 investigated the pore size distribution in biofilms generated in RBC units. Their investigations showed the presence of two different population of pores. Similar observations were also made by Okabe et al.38 They observed that two types of pores exist, macropores (20-200 µm in size) and micropores. The porosity of mature biofilms was governed by the large pores, whereas the porosity of biofilms in the early stages of development was governed by a large number of fine pores. In both types of biofilm, the porosity was found to decrease dramatically with increasing depth of the biofilm. Fu et al.39 measured the diffusivity of oxygen in biofilms using oxygen microelectrodes. They observed that the diffusivity of oxygen is related to the density of the biofilm. The diffusivity was found to decrease by almost a factor of 5 toward the bottom of the biofilm (approximately 0.5 × 10-9 m2/s) compared to that near the surface of the biofilm (approximately 2.5 × 10-9 m2/ s). Bishop et al.40 investigated the effects of biofilm structure on the transport and transformation process within the biofilm in a rotating drum bioreactor. The biofilm structure was found to be highly nonuniform with respect to depth. Along the depth, the biofilm density increases, metabolically active biomass decreases, and the porosity decreases. The viability of the biomass was found to be reduced from about 100% at the biofilm surface to 30% at the support. The porosity was found to decrease by about 25%. As a result, the effective diffusivity was found to decrease by a factor of 3-4. Beyenal et al.41 measured the oxygen diffusivity in a biofilm produced in an open-channel reactor. The diffusivity in the biofilm was found to vary by a factor of 3 over the entire thickness of the biofilm. It was also found to be dependent on the microorganisms growing within the film. The diffusivity was correlated with the biofilm density through the equation

Deff 0.43X0.92 )1Dm 11.19 + 0.27X0.99

(26)

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7. Effects of Operating Parameters on the Performance of RBCs 7.1. Effect of Speed of Rotation Friedman et al.42 investigated the effect of rotation speed on the performance of RBC system. Their experimental setup consisted of eight stages giving 37 ft2 of total surface area. They observed that the dissolved oxygen concentration decreased rapidly to a minimum value in the first or second stage and thereafter gradually increased in the subsequent stages. The rate of reduction of dissolved oxygen in the first two stages and its rate of increase in the subsequent stages were found to increase with increasing speed of rotation. Correspondingly, the extent of COD reduction was found to increase with increasing speed of rotation. 7.2. Effect of Tank Volume to Surface Area Ratio Antonie1 investigated this effect with 2- and 6-ftdiameter RBC pilot plant units. He observed that an increase in the ratio of the media surface area to the liquid volume beyond 200 m2/m3 did not increase the BOD removal rate. 7.3. Effect of Staging Antonie1 investigated the effect of the number of stages and reported that a four-stage RBC unit produced higher BOD removal rates than a two-stage unit providing the same overall surface area for the same wastewater. He also recommends that organic loading on any particular stage should not exceed 2.5 lb of soluble BOD/day/1000 ft2. 7.4. Effect of Organic and Hydraulic Loading Rates Antonie et al.22 investigated the effect of flow rate for an RBC system having a surface area of 21 000 ft2. Their results indicated that an increase in the total flow rate of wastewater causes an almost proportional reduction in the BOD removal rate and the nitrogen removal rate. However, this relationship was found to vary with the concentration or the inlet BOD level, indicating that the organic loading is also an important parameter. Stover et al.43 investigated the COD removal rate as a function of the loading rate for a six-stage RBC system. From their study, it can be seen that the extent of COD removal for all six stages could be correlated with the COD loading rates for that particular stage. Poon et al.23 investigated the performance of a fourstage RBC system. They observed that the effluent BOD concentration increased with increasing organic loading. However, this relationship was found to be different for different stages of the RBC system. Similar results were also observed for BOD removal.44 They further investigated the effect of shock loadings and observed that there is only a marginal reduction in the performance due to shock loading. Brenner et al.3 recommended that the organic loading in the first stage should be compatible with the oxygentransfer capability of the system. Overloading at any stages causes oxygen-deficient organisms to grow, resulting in poor performance. To avoid these conditions, they recommended that the first-stage organic loading should be kept in the range of 2.6-3.8 lb of BOD/day/ 1000 ft2.

Wilson and Lee45 reported that an increase in the organic loading leads to a small reduction in the removal efficiency. For example, a 20-fold increase in the organic loading rate caused only a 5% loss of removal efficiency. This could be due to the higher temperatures (equatorial climate) of the ambient air and the wastewater. Banerjee46 investigated the degradation of phenol in an RBC system and observed that an increase in the hydraulic loading leads to an increase in the phenol removal rate. Similarly, an increase in the phenol loading leads to an increase in the phenol removal rate. However, in both the cases, the conversion of phenol decreased. The rate of phenol removal was observed to be first-order with respect to the phenol concentration. 7.5. Effect of Temperature Antonie et al.22 observed that an increase in the temperature up to about 50 °F caused an increase in the BOD removal rate. Increasing temperature above this value caused only a marginal increase in the BOD removal rate. The effect of temperature was represented in terms of a temperature correction factor. The temperature correction factor was slightly higher for nitrification than for BOD removal. Banerjee46 reported that an increase in the temperature of the RBC system leads to an increase in the phenol removal rate up to a temperature of about 36 °C; above this value, the phenol removal rate is essentially constant. 7.6. Effect of Recirculation of the Effluent Klees and Silverstein47 examined the effect of recirculation of the effluent on nitrification for an RBC system. They observed that recirculation improved the nitrification for all organic loading rates. Recirculation causes dilution of the organic substrates in the RBC system. It also leads to a reduction in the available carbon, causing earlier nitrification, thereby improving the performance. They also observed that, with recirculation, the biofilm was thinner, and as a result, the biomass loading was lower, which led to a more uniform biomass loading in the various stages. However, when the recirculation ratio was increased above a certain point, the performance was found to deteriorate. Neu48 suggested the use of recirculation of the solids to the RBC system to increase the nitrification capacity of the system. Recirculation was found to increase the suspended solids present in the liquid, which also contributed to the nitrification. Wilson and Lee45 also investigated the effect of recirculation of the effluent. They observed that recirculation did not significantly affect the removal efficiency. This could be because the removal efficiency without recirculation in their work was very high (9095%). Thus, recirculation increased the total flow rate but at the same time reduced the inlet concentrations. These two factors nullified the effects of each other. 8. Mathematical Models for Performance Prediction Famularo et al.49 developed a model to predict the performance of RBC systems. The model assumes that the disks are flat and that the liquid and the biofilm developed on the disk have constant thicknesses. The liquid film mixes with the liquid in the tank instantly at the point of re-entry. The disk was divided into eight


Figure 10. Schematic representation of models formulated by Famularo et al.49 and Mueller et al.51

parts, and each part was assumed to be well-mixed. The biofilm was divided into four parts in the direction perpendicular to the disk surface. The disk surface and the biofilm are shown schematically in Figure 10. The biodegradation was assumed to occur only in the biofilm and was assumed to follow Michaelis kinetics. The rate of substrate removal was written as

RS )

µmaxX S C YX/S S + KS C + KC

(27)

The rate of uptake of oxygen was calculated as

RC )

(

)

µmaxXYC/S S C + keX YX/S S + KS C + KC

(28)

The diffusive fluxes of substrate and oxygen within the biofilm were equated with those obtained by transport from the bulk through the liquid. A material balance for the substrate was written, taking into account the convection, diffusion, and reaction in the biofilm, as

Ds

QF ∂2 S ∂S - RS ) + (SO - S) 2 Aδ ∂t ∂z

(29)

Similarly, a balance for oxygen was also written. All of the equations were solved together using finite difference methods. The kinetic parameters in the model were fitted to match the experimental measurements made on a pilot plant operated with paper mill waste and domestic wastewater. Using the kinetic parameters obtained in this manner, several predictions were made, including the effect of a step change in the concentration

of BOD in the wastewater, the effect of oxygen enrichment, the effect of media diameter, etc. The major limitations of the model are the division of the disk into only four parts in the axial direction and eight parts in the tangential direction; mixing of the liquid film with the bulk liquid; and the assumptions of perfect mixing of each of the parts, constant thicknesses of the liquid film and of the biofilm, constant biomass loading, etc. Paolini et al.50 developed a model by assuming that the oxygen does not limit the biodegradation, the liquid in the tank is perfectly mixed, concentration of the substrate in the radial and tangential directions is constant, and the liquid film thickness is constant. A mass balance equation was written for substrate in the biofilm taking into account diffusion, convection, and reaction. The flux of substrate into the biofilm was equated to its flux by mass transfer. The assumption of liquid being well-mixed in the tank is not very realistic. As before, the kinetic parameters were obtained by fitting the experimental data. Mueller et al.51 modified the earlier model of Famularo et al.49 by dividing the disk into four parts in the tangential direction and seven parts in the axial direction. Further, the overall mass balance for the substrate in the tank was written to take into account convection, transfer to biofilm, and reaction in the tank by suspended microorganisms and transfer to air. The rest of the equations were similar to those of Famularo et al.49 The model was applied for nitrification in an RBC unit. The kinetic parameters were obtained to fit the data obtained in a pilot-scale RBC unit. Chen et al.52 developed a model by considering the diffusion and reaction within the biofilm and transport from the liquid film. The model equations were similar to those developed by earlier workers. Therefore, this model also suffers from the limitations mentioned above. They applied the model to predict the simultaneous removal of organic substrate and nitrogen. Gujer and Boller53 presented a model of the RBC system taking into account biodegradation and transport processes in a manner similar to that used by the previous workers. Their major advancement over the previous work is that the variations of the biomass within the biofilm and along the reactor were considered. Within the biofilm, the mass balance for the ith dissolved species and particulate species was written as

∂2Si ∂Si ) DSi 2 - rSi ∂t ∂z

(30)

The dissolved species considered were dissolved oxygen, BOD, ammonium, nitrite, nitrate, and bicarbonate. The particulate species considered were inert material, nondegradable material, heterotrophic biomass, nitrosomonas, and nitrobacter. The rates of reaction of these species were written in terms of Michaelis kinetics. They also considered the flocculation of the microorganisms from the bulk to the biofilm and the shearing off of solids from the surface of the biofilm into the bulk liquid. Because the exact mechanisms of these processes are not known, they were considered to occur with firstorder kinetics. To relate the transport processes and biodegradation on the disk to the overall reactor performance, dissolved oxygen and soluble substrate balances were written by assuming that each stage of the RBC unit behaves as a CSTR. The attractive feature of

2046


this model is that it takes into account the possibility of biomass distribution and variations in the biomass loading in the various stages of the reactor. The biomass concentration and its distribution on the disk can be calculated by the model. Therefore, this model would have a better predictive capability than the earlier ones. However, the assumption of a well-mixed reactor is far too restrictive. This model has not been validated by direct comparison of model predictions and experimental observations. Spengel and Dzombak54 divided the disk and the biofilm into four segments in a manner similar to Famularo et al.49 The model equations are similar to those given by earlier workers (Famularo et al.49 and Mueller et al.51). Their model predictions compare well with the experimental measurements over a wide range of substrate loadings. This model suffers from the limitations that the biofilm density and thickness are assumed to be constant across the surface of the disk and the substrate utilization rates are considered to be independent of the radial location. Buchanan and Leduc55 suggested that the use of Monod kinetics for biodegradation is very complex to handle if one wishes to do optimization. They proposed the use of a variable-order model in which the rate of degradation of the substrate is assumed to be proportional to the substrate concentration raised to the power ni. The value of ni for each stage of the RBC was obtained by experiments on the laboratory scale. In fact, this approach is equivalent to representing the Monod kinetics for each stage with a power-law model. Yamaguchi et al.56 carried out experiments with an RBC system consisting of five disks. For modeling purposes, they divided the RBC system into seven parts (one between each disk) and one below the bottom edge of the disks. Each of these sections was assumed to be well-mixed. Further, the degradation process was assumed to be limited by the supply of the substrate (oxygen limitations were not considered), and Monod kinetics were assumed to be applicable. For each of the sections, a mass balance equation was written for fixed biomass, suspended biomass, substrate within the biofilm, and substrate within the liquid. The mass balance equations consisted of the following terms: (i) generation of biomass, (ii) degradation of substrate, (ii) attachment of suspended biomass, (iii) detachment of fixed biomass, (iv) death of the biomass, (v) transport of biomass and substrate from one section to another, and (vi) transport of substrate from bulk liquid to disk surface. Their model contains about 25 parameters describing transport and kinetics. The model validation was shown only for a few sets of data. Further, they did not consider the diffusion of substrate within the biofilm and the enhancement in mass transfer due to reactions occurring in the biofilm. Israni et al.57 studied the degradation of phenol in an RBC system. Their experimental setup consisted of nine disks. They modeled the system by considering that the RBC system behaves similarly to nine well-mixed tanks in series. The kinetics of degradation were assumed to be mth order with respect to oxygen and nth order with respect to the substrate. It was assumed that the rate of substrate and oxygen transfer to the biofilm were not limiting. A balance for oxygen was written within the biofilm considering its diffusion and simultaneous reaction. In this way, the enhancement in the rate of oxygen transfer due to reaction could be taken

into account. The biomass loading was assumed to be constant and no balance equation was written for it. This model was able to correlate the experimental data in their work satisfactorily. 9. Process Design Aspects The RBC process design engineer has to determine the set of parameters including the type of media, diameter of the media, speed of rotation, length of the media, number of stages, etc., so as to achieve a certain degree of treatment. Accordingly, the physical facilities, including the motor, gear system, supports, etc., also have to be designed. Moreover, the RBC configuration so determined must be such that the overall operation becomes economically viable and attractive. For wastewater treatment plants, the economic objective can be considered to be simply the minimization of the overall costs. In the early days (1970-1980), the process design criteria were largely empirical or based on general guidelines. Process design based on such general guidelines can lead to uncertainties in design. As a result, large safety margins were required. This could lead to a costly and economically unattractive design. For example, it has been suggested44,58 that process design be done on the basis of organic loading rate. However, these reports also recommend a safety factor of 1.5 for the design. Brenner et al.3 also suggested that the process design can be carried out on the basis of the organic loading rate. Tchobanglous and Burton59 also gave typical loading rates in terms of the soluble/total BOD per unit surface area of the RBC system. The problem with such oversimplified design criteria is that the intrinsic biodegradation constants and hydrodynamics of the system are not taken into consideration. Thus, for the rational design of an RBC system, it is important to consider the kinetics of the biodegradation. When mixed cultures are being used in a wastewater treatment system, the kinetics of biodegradation can change even with a small change in the wastewater composition or flow rate. This is because a change in the wastewater composition might allow a particular organism to grow at a faster rate, thereby allowing it to become a dominant microorganism, as compared to other organisms. Further, if the flow rate decreases, then the amount of substrate decreases, and the foodto-microorganism (F/M) ratio decreases, and this can lead to starvation of the microorganisms, which can lead to a change in the biodegradation kinetics. Alternatively, a reduction in the F/M ratio might lead to an increase in the availability of oxygen. This could lead to an increase in the dissolved oxygen levels, which, in turn, can change the biodegradation kinetics. The oxygen transfer from the air to the microorganisms is an important parameter that governs the dissolved oxygen levels. Thus, it is important to establish the biodegradation kinetics over a range of flow rates, dissolved oxygen concentrations, wastewater compositions, etc. The process design of an RBC system should take into consideration the underlying hydrodynamics, biodegradation kinetics, oxygen transfer, etc. Christensen et al.35 concluded that the design rules for substrate removal in biofilms used for wastewater treatment must include relations between the removal kinetics and the structure and development of the biological film. Thus, the first step in overall process design should be to carry out experiments on a laboratory-scale reactor


to determine the biodegradation kinetics over a wide range of operating conditions. These experiments can be conducted in, say, a 1-L reactor. The media to be used for the laboratory-scale experiments should be the same as that to be used in the large-scale facility. The experimental RBC system should include provisions to enable variations in the liquid submergence, speed of rotation, flow rate and concentration of the substrate, etc. Experiments should be carried out to investigate the effects of hydraulic retention time. At each retention time, experiments should be carried out to investigate the effects of speed of rotation, diameter, and inlet BOD concentration. For a particular retention time, an increase in the speed of rotation increases the rate of mass transfer from the gas phase, which leads to an increase in dissolved oxygen in the liquid phase. This enables quantification of the effect of the dissolved oxygen level on the kinetics of biodegradation. An increase in the speed of rotation also increases the shear levels within the reactor. This might cause more detachment of the microorganisms. The level of suspended solids can therefore increase. Gujer and Boller53 assumed that the rate of detachment of microorganisms is first-order with respect to shear. Experiments at various speeds of rotation will enable quantification of the effect of shear. At a particular retention time, an increase in the inlet substrate concentration leads to an increase in the F/M ratio. These experiments then quantify the effect of the F/M ratio on the kinetics of biodegradation. For each experiment, the extent of BOD/ COD removal should be measured. This will enable computation of the rate of BOD/COD removal per unit volume or per unit surface area of the RBC system, which can be calculated as

rV )

(∆BOD)L (∆BOD)L or rA ) VL A

(31)

Further, the dissolved oxygen levels in each experiment should also be measured. It is desirable to estimate the biomass loading on the media in each of the experiments. This will enable quantification of the effects of biomass loading. The kinetics of biodegradation can then be established using either a power-law model or Monod-type kinetics. For the power-law model, the rates can be fitted with the equation

rV or rA ) k1Xk2(C)k3Sk4

(32)

The values of the constants k1-k4 can be fitted to match the experimental data. The Monod-type kinetics can be represented by eqs 27 and 28. Once the biodegradation kinetics have been measured over a wide range of parameters, they are expected to remain unchanged upon scale-up. With these kinetics, it is possible to design the largescale RBC system. The following stepwise procedure is recommended for the process design: Step 1. Assume a certain diameter of the RBC media. Step 2. Assume a certain speed of rotation. Step 3. Calculate the tip speed and estimate the shear levels and the biomass loading that would be expected in the reactor. Step 4. Calculate the value of the mass-transfer coefficient. Step 5. Calculate the dissolved oxygen level such that the rate of oxygen uptake due to biodegradation equals the rate of mass transfer.

Step 6. Knowing the rate of biodegradation and the desired degree of removal, calculate the number of disks/ surface area of media required. Step 7. Because the processes of biodegradation, generation of cell mass, consumption of oxygen and substrate are interlinked, the above steps will have to be carried out iteratively. Repeat this procedure for various speeds, repeating steps 3-6. Step 8. This process will finally generate a set of speeds and the corresponding surface areas of media required to achieve a certain degree of BOD removal. Step 9. The fixed costs of the system depend on the surface area of the media, and the operating costs depend on the speed of rotation. Determine the annualized cost as a certain combination of the fixed costs plus the operating costs. Determine the annualized cost for the various speeds obtained from step 8. Choose the speed of rotation at which the annualized cost is a minimum. 10. Comparison with Other Wastewater Treatment Systems Secondary wastewater treatment systems employ microorganisms for the degradation of the organic waste. In such systems, a variety of equipment is used for the purpose of aeration such as surface aerators, diffusers, RBCs, trickling filters, etc. The rate of degradation of the waste depends on the rate at which the microorganisms metabolize the waste and the loading of biomass. The rate of metabolism depends on the state of the microorganisms, the availability of the substrate and oxygen, etc. The biomass loadings are generally much higher for trickling filters and RBCs as compared to the activated sludge process. This is mainly because the former employ fixed biological films, whereas the latter employs suspended solids. For a given wastewater flow, inlet BOD, and desired degree of BOD removal, the availability of substrate would be same in all of the equipment. The choice between the various equipment would then boil down to the ability of the equipment to supply oxygen efficiently. This is frequently characterized by the oxygen-transfer efficiency, which is the amount of oxygen transferred per unit energy consumed (in units of kilograms of O2 per kilowatt-hour). Thus it was thought desirable to compare the oxygen-transfer efficiency for the various types of equipment mentioned above. For the purpose of illustration, assume that the wastewater flow rate is 30 000 m3/day, with an inlet BOD level of 1000 mg/L and a desired outlet level of 100 mg/L. This would mean that the oxygen-transfer requirement, denoted by M, is 27 000 kg/day. The oxygen-transfer efficiencies of various types of equipment in achieving this overall oxygen-transfer rate are compared below. 10.1. Surface Aerators Tchobanglous and Burton59 reported that the oxygentransfer efficiency of various types of mechanical aerators is in the range of 1.2-3.0 kg of O2/kWh. McWhirter et al.60 analyzed the performance of two working aerators and observed that the oxygen-transfer efficiency is about 1.8 kg of O2/kWh. Similar efficiencies were also observed for diffuser aeration systems. To calculate the oxygen-transfer efficiency, the following procedure is adopted:

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(i) A reactor size is chosen (VSA). (ii) The volumetric rate of oxygen transfer is calculated as M/VSA. (iii) Using the saturation concentration of oxygen (C*) as 8 mg/L, the desired mass-transfer coefficient is calculated as kLa ) M/(VSAC*). This assumes that the dissolved oxygen concentration is 0. Although this is undesirable, it gives only the minimum possible value of kLa that is acceptable. In practice, it would be necessary to have kLa higher by 10-20%, so as to achieve a dissolved oxygen concentration of around 2 mg/L. (iv) The power required to achieve this mass-transfer coefficient is calculated using the correlations recommended earlier.10 For this purpose, a value for the power consumption per unit mass is assumed. With this value, the gas hold-up and the mean bubble size are calculated. These values are used to calculate the interfacial area and finally the mass-transfer coefficient, kLa. The power consumption per unit volume is varied so that the calculated kLa matches the desired value obtained in iii. (v) This approach allows for the calculation of the oxygen-transfer efficiency as a function of the power consumed per unit volume of the reactor. (vi) The fixed cost of the activated sludge process is calculated in the following manner: the costs of the aeration tank and the secondary clarifier are estimated from vendors’ quotations as $550 × (tank volume)0.7, $700 × (clarifier volume)0.8, and $25 000 per 100 hp surface aerator. The operating cost is calculated assuming 8000 working hours per year and an electricity cost of $0.035/kWh. The annualized cost is computed as 20% of the fixed cost plus the operating cost. 10.2. RBC Systems The following procedure is adopted for estimation of the oxygen-transfer efficiency of an RBC: (i) The diameter of the disk is assumed to be 4 m (similar to standard or high-density media having 12ft diameters). (ii) A value for the speed of rotation is assumed. (iii) The kL value is calculated using the correlation of Boumansour and Vasel.30 (iv) Because the oxygen requirement M is known, the surface area can be calculated. (v) Assuming that each unit consists of a 27-ft-long shaft and 150 000 ft2 area per unit, the number of units required can be calculated. The power requirements can be calculated assuming that the power number per disk is 0.1. This enables calculation of the oxygen-transfer efficiency. (vi) The above procedure can be repeated for various speeds of rotation to compute the oxygen-transfer efficiency as a function of the power consumption per unit reactor volume. (vii) The fixed cost of an RBC system depends on the total surface area of the media. The weight of polyethylene that would be required is calculated by assuming that the density of polyethylene is 900 kg/m3. To account for the fabrication losses, 10% extra polyethylene is assumed. The fixed cost of the RBC system is calculated assuming that the cost of polyethylene is about $0.5/ kg. As before, the operating cost is calculated assuming 8000 working hours per year and an electricity cost of

Figure 11. Comparison of the oxygen-transfer efficiencies of an RBC system, the activated sludge process, and trickling filters: ) RBC system; s, activated sludge process; ‚ ‚ ‚, trickling filters.

$0.035/kWh. The annualized cost is computed as 20% of the fixed cost plus the operating cost. 10.3. Trickling Filters For the process design of trickling filters, it was assumed that the filter was operated with the recycle ratio of 2:1 and the air flow was considered to be 5 times the stoichiometric requirement. This ensures that oxygen depletion does not lead to anaerobic conditions deep within the filter. The trickling filter is assumed to be filled with rocks. The following procedure is adopted: (i) The total liquid flow is first calculated as the sum of the wastewater flow plus the recirculation flow. (ii) A certain value of the superficial liquid velocity is assumed. Because the total flow rate is known, the desired cross-sectional area of the trickling filter can be calculated. (iii) Using the above results, the mass-transfer coefficient and pressure drops are calculated with correlations given by Treybal.61 (iv) Because the total oxygen-transfer requirements are known, the packed volume can be calculated, and hence, the height can be determined. In practice, if the height is very large, it might be necessary to split the operation into two or more trickling filter units in series/ parallel. (v) Knowing the pressure drops and the flow rates of air and wastewater, the power requirements for the wastewater pump and the air blower are calculated. This allows for the calculation of the oxygen-transfer efficiency. The above procedure is repeated for several superficial liquid velocities, and the oxygen-transfer efficiency as a function of power input per unit volume is plotted. (vi) The fixed cost of the trickling filter is calculated by assuming the cost of the media to be $500/m3 of packings. The operating cost is calculated assuming 8000 working hours per year and an electricity cost of $0.035/kWh. The annualized cost is computed as 20% of the fixed cost plus the operating cost. The above three types of equipment are compared on the basis of their oxygen-transfer efficiencies and their annualized costs of operation. The oxygen-transfer efficiencies are plotted in Figure 11 for all three systems. From the graph, it can be seen that, at low power consumption, the oxygen-transfer efficiency of the RBC system is much higher than those of the other two systems. However, the RBC oxygen-transfer efficiency decreases rapidly with increasing power consumption per unit volume. At high power consumption per unit


Figure 12. Comparison of the annualized costs of operation of an RBC system, the activated sludge process, and trickling filters: ) RBC system; s, activated sludge process; ‚ ‚ ‚, trickling filters.

volume, the trickling filters might be most efficient. Figure 12 compares the annualized costs of operation for these three types of equipment. From this figure, it can be seen that the optimum costs for RBC and the activated sludge process occur at low power consumption per unit volume. For RBC systems, the optimum cost occurs at speeds of 3-4 rpm, which are slightly higher than those used in practice. For the activated sludge process, the optimum cost corresponds to a biomass loading of about 2000 mg/L. This is the level of biomass loading normally kept in the activated sludge process. For the trickling filter process, because of the high cost of packing material, the operation becomes more economical at high power consumptions per unit volume. It might not be feasible to operate at such high power consumptions because of excessive slugging of microorganisms from packings. A comparison of the three costs reveals that the optimum costs of the activated sludge process and trickling filters would work out to be approximately the same. The annualized cost of the RBC system appears to be on the higher side. It must be kept in mind that these calculations are based only on oxygen-transfer considerations. RBC systems and trickling filters facilitate high biomass loadings and plug-flow behavior, which will reduce the volume requirements, and hence, can considerably reduce the fixed costs. Further, the oxygen-transfer rate is enhanced as a result of reactions occurring within the biofilm. 11. Conclusions Rotating biological contactors (RBCs) offer an alternative treatment technology to the conventional activated sludge process. The principal advantages of RBCs stem from the fact that the interfacial area generated is very high and practically independent of the speed of rotation, unlike in the activated sludge process. Generally, RBC media are available are in the form of corrugated plastic sheets. The interfacial area provided by them is on the order of 150-250 m2/m3 of liquid. The flow pattern with plain flat disks can be characterized as plug flow with axial dispersion. The liquid volume entrained on the disk surface can be estimated from the correlation given earlier by Vaidya and Pangarkar.16 The literature suggests that the power consumption of an RBC unit can be characterized in terms of a power number. Judging from the literature, the power number for a single disk is in the range 0.05-0.2, typically 0.1. The mass-transfer aspects have been studied for a set

of plain disks, and different models have been proposed. The mass-transfer coefficients can be estimated using the correlation of Boumansour and Vasel.30 A number of mathematical models have been proposed for estimation of the performance of RBC systems. These models take into account the transport and biochemical reactions occurring in the liquid and biofilm. Of these, the most detailed so far is the model presented by Gujer and Boller.53 For process design, it is highly desirable to carry out laboratory/pilot-scale experiments. A stepwise process design algorithm has been presented. It is observed that an RBC system shows much higher oxygen-transfer efficiency than the conventional equipment such as activated sludge processes employing surface aerators or diffuser aeration and trickling filters. 12. Suggestions for Future Work This literature review has highlighted various aspects that should be investigated in the future. The various types of media should be studied in detail. For each type of media, the power consumption, hydrodynamics, and mass-transfer characteristics need to be investigated. The media can then be compared on a uniform basis, such as equal power consumption, for rational selection of the media. The biomass loading on the media is determined by the rate of growth of the microorganisms, their ability to adhere to each other and to the media and the shear level they experience. These parameters also need to be quantified for a variety of media. The hydrodynamics of overall RBC systems need to be evaluated to ascertain whether such systems behave similarly to plug-flow reactors or CSTRs. The results of such studies will have an important bearing on scaleup. 13. Nomenclature A ) surface area, m2 Ad ) wetted surface area per disk, m2 a ) surface/interfacial area per unit volume, m2/m3 BOD ) biochemical oxygen demand, mg/L C ) concentration of dissolved oxygen, kmol/m3 Ca ) capillary number, µvc/σ Ci ) initial concentration of dissolved oxygen, kmol/m3 C* ) saturation concentration of dissolved oxygen, kmol/m3 COD ) chemical oxygen demand, mg/L CSTR ) continuous stirred-tank reactor D ) diameter of the disk, m DL ) axial dispersion coefficient, m2/s Dm ) molecular diffusivity of oxygen, m2/s DS ) molecular diffusivity of substrate, m2/s d ) disk spacing, m dh ) depth of submergence of the disk, m Fr ) Froude number g ) acceleration due to gravity, m/s2 gpd ) gallons per day H ) height of submergence of the disk from the bottom, m HL ) hydraulic loading, m3/(m2 s) k1, k2, k3, k4 ) constants in eq 32 KC, KS ) constants in eq 27 ke ) death rate constant, s-1 kLa ) volumetric mass-transfer coefficient, s-1 kL ) mass-transfer coefficient for the liquid film exposed to air, m/s kL′ ) mass-transfer coefficient for the liquid in the boundary layer flow, m/s L ) length of the reactor, m

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Mn ) COD removal rate N ) speed of rotation, s-1 NC ) flux of dissolved oxygen, kmol/(m2 s) NP ) power number P ) power consumption, W Q ) flow rate of liquid entrained on the disk surface, m3/s r ) radial coordinate, m R ) radius of the disk, m RC ) rate of oxygen absorption, kmol/s RS ) rate of substrate removal, kmol/(m3 s) rA ) rate of substrate removal per unit area, kmol/(m2 s) rV ) rate of substrate removal per unit volume, kmol/ (m3 s) RBC ) rotating biological contactor RDC ) rotating disk contactor Re ) Reynolds number RT ) retention/residence time, s S ) concentration of the substrate, kmol/m3 Si ) concentration of the ith, substrate, kmol/m3 SO ) initial concentration of the substrate, kmol/m3 Sh ) Sherwood number TO ) dimensionless film thickness, δ(Fg/µvc)1/2 t ) time, s tC ) contact time, s U ) velocity of liquid feed, m/s V ) volume of liquid entrained on the disk surface, m3 VL ) volume of liquid in the tank, m3 vc ) tip velocity of the disk, m/s vr ) velocity in the radial direction, m/s vθ ) velocity in the tangential direction, m/s X ) biomass loading mg/L x ) axial coordinate in the Cartesian system, m YC/S ) yield coefficient YX/S ) yield coefficient Z ) maximum film thickness, m z ) axial coordinate in the polar system of coordinates, m Greek Letters R ) constant relating power number to Reynolds number δ ) film thickness, m δBL ) boundary layer thickness, m ) fractional voidage in the reactor θ ) tangential coordinate λ1, λ2 ) constants with units as given in the text µ ) viscosity, Pa s µmax ) maximum growth rate, s-1 ν ) kinematic viscosity, m2/s FL ) density of liquid, kg/m3 τrz ) shear stress in the r-z direction, N/m2 τzθ ) shear stress in the z-θ direction, N/m2 φ ) angle ω ) angular speed of rotation, s-1

14. Literature Cited (1) Antonie, R. L. Fixed Biological Surfaces Wastewater Treatment; CRC Press: Cleveland, OH, 1976. (2) Greaves, F. E.; Thorp, B.; Critchley, R. F. Operational Performance of Package Sewage Treatment Plants in North West England. Water Sci. Technol. 1990, 22, 25-32. (3) Brenner, R. C.; Heideman, J. A.; Opatken, E. J.; Petrasek, A. C. Design Information on Rotating Biological Contactors; U.S. EPA Report EPA-600/2-84-106; U.S. Environmental Protection Agency: Washington, DC, 1984. (4) Mba, D.; Bannister, R. H.; Findlay, G. E. Mechanical Redesign of the Rotating Biological Contactor. Water Res. 1999, 33, 3679-3688. (5) Doraiswamy, L. K.; Sharma, M. M. A. Heterogeneous Reactions: Analysis, Examples, and Reactor Design; Wiley-Interscience: New York, 1984. (6) Wanner, J.; Sykora, M.; Kos, M.; Miklenda, J.; Grau, P. Packed Cage RBC with Combined Cultivation of Suspended and Fixed Film Biomass. Water Sci. Technol. 1990, 22, 101-111.

(7) Ware, A. J.; Pescod, M. B.; Storch, B. Evaluation of Alternatives to Conventional Disc Support Media for Rotating Biological Contactors. Water Sci. Technol. 1990, 22, 113-117. (8) McManus, M. J. Media for Rotating Biological Contactor. U.S. Patent 5,401,398, 1995. (9) Nahid, P.; Vossoughi, M.; Alemzadeh, I. Treatment of Bakers Yeast Wastewater with a Biopack System. Process Biochem. 2001, 37, 447-451. (10) Patwardhan, A. W.; Joshi, J. B. Design of Gas Inducing Reactors. Ind. Eng. Chem. Res. 1999, 38, 49-80. (11) Gai, S.; Kruger, K.; Kanne, L.; Mohr, K. The Rotary Trickle-Bed ReactorsA New Reactor Concept for Biological Gas Purification. Eng. Life Sci. 2001, 1, 5-14. (12) Yamane, T.; Yoshida, F. Absorption in a rotating disk gasliquid contactor. J. Chem. Eng. Jpn. 1972, 5, 55-59. (13) Bintanja, H. H. J.; van der Erve, J. J. V. M.; Boelhouwer, C. Oxygen Transfer in a Rotating Disc Treatment Plant. Water Res. 1975, 9, 1147. (14) Zeevalkink, J. A.; Kelderman, P.; Boelhouwer, C. Liquid film thickness in a rotating disc gas-liquid contactor. Water Res. 1978, 12, 577-581. (15) Suga, K.; Boongorsrang, A. A new model of mass transfer in a rotating disk contactor. Chem. Eng. Sci. 1984, 39, 767-773. (16) Vaidya, R. N.; Pangarkar, V. G. Hydrodynamics and mass transfer in a rotating biological contactor. Chem Eng. Commun. 1985, 39, 337-354. (17) White D. A.; Tallmadge J. A. Theory of drag out of liquids on flat plates. Chem. Eng. Sci. 1965, 20, 33. (18) Vaidya, R. N.; Pangarkar, V. G. Convective Diffusion Model for Mass Transfer in a Rotating Biological Contactor: Disc Submergence

Rotating Biological Contactors: A Review - American Chemical Society

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