NOTES ON LESSON-BT 2353-III YEAR A&BDEPARTMENT OF BIOTECHNOLOGY

RAJALAKSHMI ENGINEERING COLLEGE,THANDALAMCHENNAI

Prepared by

Kavitha VijayaraghavanLecturer, Department of Biotechnology,

Rajalakshmi Engineering College,Thandalam

BT 2353 BIOPROCESS ENGINEERING

AIMThis course aims to develop the skills of the students in the area of BioprocessEngineering. This will be a pre-requisite for a few elective courses and for project inBioprocess Technology.OBJECTIVES At the end of the course, the student would have learnt about stirred Tankreactors and configuration of various reaches, and how to model and similar aBio process. This will help the student to undertake project in the area of Bioprocess Technology.UNIT I ANALYSIS OF STR 8Stirred tank reactor - non-ideality, RTD and stability analysis, tanks in series anddispersion models – application to design of continuous sterilizer.UNIT II ANALYSIS OF OTHER CONFIGURATIONS 8Packed bed reactor, airlift reactor, fluidized bed reactor bubble column reactors – nonideality,RTD and stability analysis.UNIT III BIOREACTOR SCALE – UP 9Regime analysis of bioreactor processes, oxygen mass transfer in bioreactors -microbial oxygen demands; methods for the determination of mass transfer coefficients;mass transfer correlations. Scale up criteria for bioreactors based on oxygen transfer,power consumption and impeller tip speed.

UNIT IV MODELLING AND SIMULATION OF BIOPROCESSES 12Study of structured models for analysis of various bioprocess – compartmental models,models of cellular energetics and metabolism, single cell models, plasmid replication andplasmid stability model. Dynamic simulation of batch, fed batch, steady and transientculture metabolism.UNIT V BIOREACTOR CONSIDERATION IN ENZYME SYSTEMS 8Analysis of film and pore diffusion effects on kinetics of immobilized enzyme reactions;formulation of dimensionless groups and calculation of effectiveness factors. Design ofimmobilized enzyme reactors – packed bed, fluidized bed and membrane reactors.TOTAL : 45 PERIODSTEXT BOOKS1. Anton Moser, “Bioprocess Technology”, Kinetics and Reactors”, Springer Verlag.2. James E. Bailey & David F. Ollis, “Biochemical Engineering Fundamentals”,McGraw-Hill.3. Shuler and Kargl,Bioprocess Engineering, Prentice Hall , 1992.REFERENCES1. James M. Lee, “Biochemical Engineering”, PHI, USA.2. EMT.EL-Mansi.CFA.Bryce, A.L.Demain, AR.Allman: Fermentation Microbiology andBiotechnology, Second Edition 2007.3. Harvey W. Blanch, Douglas S. Clark, “Biochemical Engineering”, Marcel Decker Inc.

UNIT I ANALYSIS OF STR 8Stirred tank reactor - non-ideality, RTD and stability analysis, tanks in series anddispersion models – application to design of continuous sterilizer.Continuous Stirred Tank Reactors (CSTRs)

Type of Reactor

Characteristics

Continuously Stirred Tank Reactor (CSTR)

Run at steady state with continuous flow of reactants and products; the feed assumes a uniform composition throughout the reactor, exit stream has the same composition as in the tank

Kinds of Phases Present

Usage Advantages Disadvantages

1. Liquid phase

2. Gas-liquid rxns

3. Solid-liquid rxns

1. When agitation is required

2. Series configurations for different concentration streams

1. Continuous operation

2. Good temperature control

3. Easily adapts to two phase runs

4. Good control

5. Simplicity of construction

6 Low

1. Lowest conversion per unit volume

2. By-passing and channeling possible with poor agitation

operating (labor) cost

7. Easy to clean

General Mole Balance Equation

Assumptions

1) Steady state therefore

2) Well mixed therefore rA is the same throughout the reactor

Rearranging the generation

In terms of conversion

Reactor Sizing

Given –rA as a function of conversion, –rA = f(X), one can size any type of reactor. The volume of a CSTR can be represented as the shaded areas in the Levenspiel Plot shown below:

Reactors in Series

Given –rA as a function of conversion, , –rA = f(X), one can also design any sequence of reactors in series provided there are no side streams by defining the overall conversion at any point.

Mole Balance on Reactor 1

Mole Balance on Reactor 2

Given –rA = f(X) the Levenspiel Plot can be used to find the reactor volume

For a PFR between two CSTRs

Problems with a straight-forward calculation.

The following reaction takes place in a CSTR:

 

Pure A is fed to the reactor under the following conditions:

FAo = 10 mol/minCAo = 2 mol/dm3

V= 500 dm3 and k=0.1/min

X=?

Rate Law: -rA = kCA

What is the conversion in the CSTR?

 

The design equation for a CSTR is:

Derive Equation

 

Because the reaction is elementary, the combined mole balance and rate law becomes:

 

From stoichiometry, this equation becomes:

 

 

Solving for X, we get the following equation:

 

then we substitute numerical values for our variables and calculate a conversion of:

 

Problems that require intermediate calculations or manipulations.

The following reaction takes place in a CSTR:

 

Pure A is fed to the reactor under the following conditions:

FAo = 10 mol/min

CAo = 2 mol/dm3

At T=350 K

V= ? and k=?

X=0.75

Rate Law: -rA = kCA

The activation energy: E=20 kcal/mol

 

What is the conversion in a PFR at 325 K with the same volume?

Problems that require intermediate calculations or manipulations.

In this problem we have to find a way to relate these two reactors. At first, this question looks unsolvable, but then we recall the following relationship between k1 and k2:

Derive Equation

 

In the Homogeneous Example 1 Solution, we derived the following equation for a CSTR:

 

Our next step is to group our unknowns on one side of the equation and solve for their product:

 

Then, we use our relationship between k1 and k2:

 

k2V=3.3

 

Next, we look at our PFR design equation:

Derive Equation

 

We integrate this equation over our limits:

 

and solve for X:

 

Recalling our result for k2V that we determined above, we get:

 

X=0.72

Problems that are over-specified.

The following irreversible reaction takes place in a CSTR:

 

It takes place under the following conditions:

FTo = 40 mol/min

CAo = 2 mol/dm3

The feed is 75 mol% A

& 25 mol% inerts.

X=?

 

k(400 K) = 0.1/min

V = 500 dm3 and T = 400 K

 

This reaction follows the first order rate law: -rA = kCA

The activation energy (E) is 10 kcal/mol & the Arrhenius constant (A) is 3*104/min

 

What is the conversion?

Problems that are over-specified.

This problem might confuse the student with excess information about the rate constant, but there is a straight-forward solution.

We begin with the design equation for a CSTR is:

Derive Equation

 

Following the calculations in the Homogeneous Example 1 Solution, we get the equation for the conversion:

 

The problem gave us every variable except for FAo. FAo is just the mole fraction of A (xA) multiplied by the total initial flow rate (FTo).

 

Then, we substitute for our variables and calculate a conversion of:

 

Stirred Tank Reactor

        A batch stirred tank reactor is the simplest type of reactor.  It is composed of a reactor and a mixer such as a stirrer, a turbine wing or a propeller. The batch stirred tank reactor is illustrated below:

        This reactor is useful for substrate solutions of high viscosity and for immobilized enzymes with relatively low activity. However, a problem that arises is that an

immobilized enzyme tends to decompose upon physical stirring. The batch system is generally suitable for the production of rather small amounts of chemicals.

        A continuous stirred tank reactor is shown below:

        The continuous stirred tank reactor is more efficient than a batch stirred tank reactor but the  equipment is slightly more complicated.

Continuous Sterilization

Batch sterilization wastes energy and can overcook the mediumBatch sterilization uses steam or direct firing to elevate the temperature, and then cooling water stops the process and brings the material back toward room temperature. Both the heat and the cooling water are spent with no opportunity for energy recovery. Large volumes should be passed continuously through heat exchangers for energy economy with the hot, treated fluid heating the cold, incoming feed.

One method of continuous sterilization injects steam into the medium (no heat exchanger). The medium stays in a loop for a predetermined holding time until the entire medium is sterile.

Better heat economy comes from substituting heat exchangers for direct steam injection. Instead of having a cold water stream to cool the sterile media, the lower temperature unsterile media stream absorbs heat from the warm stream, cooling the sterile media.

A system for continuous sterilization has a holding coil for detention long enough to kill all of the microorganisms. The medium from a make up vessel flows through the exchanger, is held in the coil, and passes back through the heat exchanger, heating more unsterile medium while becoming cool itself, as it is collected in a sterile fermenter.

This design would work only with an exchanger with infinite heat transfer area because there is no driving force for heat transfer as the temperatures for the two streams approach closely. A real design would have another small exchanger to raise the

temperature to the setpoint after the main exchanger has done all it can do. There is no need for a cooler before entering the fermenter because it has a jacket or coils for temperature control that can easily handle this load.

Heat economy is not important for a small pilot plant unit for continuous sterilization, so direct steam injection is simpler. A heat exchanger is then needed with cooling water to bring the medium back down quickly to a temperature at which it is not over cooked. 

Advantages: Uniform steam requirements throughout the duration of the sterilization Simplified process control Shorter sterilization time means less thermal degradation of medium

Disadvantages: High demand for steam in a shorter period of time than batch Concentration of media becomes dilute due to steam condensation Since steam is actually dispersed in media, steam must be clean to avoid contamination

High temperatures for short times are used in preparing nutrient media for industrial fermentations and in pasteurizing milk, because this causes less damage to biochemicals than more prolonged times at lower temperatures. This exploits the temperature effects on activation energies because bacterial killing is affected by a temperature change more than is heat destruction of biochemicals.

Shell and Tube Exchangers

When the flow in a heat exchanger is countercurrent, the outlet temperature of the stream being heated can approach the temperature of the hot stream to be cooled. There is an attempt to show this in the sketch. There are gradients on the shell side as well.

The Shell and tube exchanger is not as well-suited to continuous sterilization as the plate-and-frame type of exchanger.

Sandwich of Plates. This sketch did not turn out well.  Much better sketch by Lawrence Bernstone, 1995 The idea is that the countercurrent flow in alternate sections gives a gradient from coolest to hottest in each plate of the sandwich. In chemical processing, a packed bed is a hollow tube, pipe, or other vessel that is filled with a packing material. The packing can be randomly filled with small objects like Raschig rings or else it can be a specifically designed structured packing.

UNIT II ANALYSIS OF OTHER CONFIGURATIONS 8Packed bed reactor, airlift reactor, fluidized bed reactor bubble column reactors – nonideality,RTD and stability analysis

Packed bed reactor

The purpose of a packed bed is typically to improve contact between two phases in a chemical or similar process. Packed beds can be used in a chemical reactor, a distillation process, or a scrubber, but packed beds have also been used to store heat in chemical plants. In this case, hot gases are allowed to escape through a vessel that is packed with a refractory material until the packing is hot. Air or other cool gas is then fed back to the plant through the hot bed, thereby pre-heating the air or gas feed.

Applications

In industry, a packed column is a type of packed bed used to perform separation processes, such as absorption, stripping, and distillation. A packed column is a pressure vessel that has a packed section.[1] The column can be filled with random dumped packing or structured packing sections, which are arranged or stacked. In the column, liquids tend to wet the surface of the packing and the vapors pass across this wetted surface, where mass transfer takes place. Packing material can be used instead of trays to improve separation in distillation columns. Packing offers the advantage of a lower pressure drop across the column (when compared to plates or trays), which is beneficial while operating under vacuum. Differently shaped packing materials have different surface areas and void space between the packing. Both of these factors affect packing performance.

Another factor in performance, in addition to the packing shape and surface area, is the liquid and vapor distribution that enters the packed bed. The number of theoretical stages required to make a given separation is calculated using a specific vapor to liquid ratio. If the liquid and vapor are not evenly distributed across the superficial tower area as it enters the packed bed, the liquid to vapor ratio will not be correct and the required separation will not be achieved. The packing will appear to not be working properly. The height equivalent to a theoretical plate (HETP) will be greater than expected. The problem is not the packing itself but the mal-distribution of the fluids entering the packed bed. These columns can contain liquid distributors and redistributors which help to distribute the liquid evenly over a section of packing, increasing the efficiency of the mass transfer.[1] The design of the liquid distributors used to introduce the feed and reflux to a packed bed is critical to making the packing perform at maximum efficiency.

Packed columns have a continuous vapor-equilibrium curve, unlike conventional tray distillation in which every tray represents a separate point of vapor-liquid equilibrium. However, when modeling packed columns it is useful to compute a number of theoretical plates to denote the separation efficiency of the packed column with respect to more traditional trays. In design, the number of necessary theoretical equilibrium stages is first determined and then the packing height equivalent to a theoretical equilibrium stage, known as the height equivalent to a theoretical plate (HETP), is also determined. The total packing height required is the number theoretical stages multiplied by the HETP.

Columns used in certain types of chromatography consisting of a tube filled with packing material can also be called packed columns and their structure has similarities to packed beds.

Packed bed reactors

Packed bed reactors can be used in chemical reaction. These reactors are tubular and are filled with solid catalyst particles, most often used to catalyze gas reactions. [2] The chemical reaction takes place on the surface of the catalyst. The advantage of using a packed bed reactor is the higher conversion per weight of catalyst than other catalytic reactors. The reaction rate is based on the amount of the solid catalyst rather than the volume of the reactor

THREE PHASE FLUIDIZED BED REACTOR A method of suspending solid particles is to fluidize the bed using upward flow

of the liquid.

If gas is introduced at the bottom of the bed three phase contacting is achieved.

As in bubble columns,agitated tank contacters hydrodynamic interaction between the bubbles and particles occur.

Large particles can cause break up of large bubbles. At present time 3 phase fluidised bed are not often chosen for gas – liquid –solid

reactions despite their advantages of good heat and mass transfer,the reason being development of three phase fluidised bed reactor for the new process would necessarily require extensive experimental work.

Also scale up from pilot to large scale is uncertain due to insufficient knowledge.

In contrast the scale up of trickle bed reactors is easy and might be chosen for a process inspite of the advantages of the 3 phase fluidized bed reactor.

The 3 phase fluidized bed reactors so far developed commercially for processes such as hydrogenation of petroleum residues using bed of particles of various sizes.

Reactors with small molecules tend to produce large bubbles.These are similar to bubble column

slurry reactors. A three phase fluidised bed using large particles requires high liquid flow rates in

order to maintain the particles in the fluidised state but have the advantage of producing smaller bubbles and hence

large interfacial area for mass transfer. One of the areas of uncertainity in 3 phase design is the mixing that occurs in gas

and liquid phases. As particle size increases flow approximates to plug flow The overall rate of the reaction calculated from the 3 phase fluidised bed reactor is

approximately one tenth of the rate calculated for agitated slurry reactor. The main reasons are very poor effectiveness factor,

relatively smaller external surface area for mass transfer caused by using larger particles.

Although the three phase reactor does not appear to perform favorably compared to agitated tank sluury

reactor it should be remembered that fluidised bed does not have a sealing or other mechanical complications associated with agitators.

RESISTANCES The various resistances active in a 3 phase fluidised bed reactor are

1. Gas liquid mass transfer resistance ( 1/kla).2. Liquid solid mass transfer resistance ( 1/ksapεp).3. Reaction within particle ( 1/k1ηεp).

The overall rate Rt = 1/ (1/kla+ 1/ksapεp+ 1/k1ηεp)

A fluidized-bed reactor is a combination of the two most common, packed-bed and stirred tank, continuous flow reactors. It is very important to chemical engineering because of its excellent heat and mass transfer characteristics. The fluidized-bed reactor can be seen below:

        In a fluidized-bed reactor, the substrate is passed upward through the immobilized enzyme bed at a high enough velocity to lift the particles. However, the velocity must not be so high that the enzymes are swept away from the reactor entirely. This causes some mixing, more than the piston-flow model in the packed-bed reactor, but complete mixing as in the CSTR model. This type of reactor is ideal for highly exothermic reactions because it eliminates local hot-spots, due to its mass and heat transfer characteristics mentioned before. It is most often applied in immobilized-enzyme catalysis where viscous, particulate substrates are to be handled.

FLUIDIZED BED REACTOR

The first commercial use of a fluidized-bed reactor, in the 1920s was for the gasification of coal to supply CO and H, for the production of synthetic chemicals.

The technology has been developed also for a number of other processes, including other catalytic reactions for the production of chemicals and synthetic gasoline by the Fischer-Tropsch process, and noncatalytic processes for the roasting of sulfide ores, and the incineration of solid waste.

Fluidized-bed and similar technology has also been used for nonchemical operations in heat exchange, drying of solids, and coating processes analogous to chemical vapor deposition.

Consider a bed of solid particles initially fixed in a vessel, and the upward flow through the bed of a fluid introduced at many points below the bed, as indicated schematically . The rate of flow of fluid is characterized by the superficial linear velocity us, that is, the velocity calculated as though the vessel were empty. At relatively low values of us, the bed remains a fixed bed of particles, although it expands somewhat as us, increases.

As u, increases further, a point is reached at which the bed begins to “lift”. Beyond this point, over a range of increasing us, the action of the multiphase fluid + solid region (bed) may resemble that of a vigorously boiling liquid with an apparent upper surface from which rising fluid disengages.

This action is referred to in general as fluidization of the bed (in the case above, a “bubbling fluidized bed”.

As U, increases still further, solid particles become elutriated from the fluidized region by entrainment in the rising fluid.

At sufficiently high values of us, the entire bed of particles may be entrained with the fluid and carried overhead out of the vessel. Other types of intermediate behavior occur, usually undesirable, and we have simplified the picture considerably.

When a chemical reaction occurs in the system, each of these types of behavior gives rise to a corresponding type of reactor. These range from a fixed-bed reactor, to a fluidized-bed reactor without significant carryover of solid particles, to a fast-fluidized-bed reactor with significant carryover of particles, and

ultimately a pneumatic-transport or transport-riser reactor in which solid particles are completely entrained in the rising fluid. (BFB)

(Pneumatic Conveying) (Fast Fluidized)

At higher gas velocities the BFB transforms into TB-no distinct bubbles,much churning and violent solid movement.The surface of the dense bed fades and solids are increasingly found in the lean region above the bed.The flow of gas is between PF and BFB.The concentration of solids in the upper lean region can be represented by exponential decay function.

In a fast-fluidized bed, the fluidization velocity is very high, resulting in significant entrainment of solid particles. Consequently, there is a concentration gradient of solid particles through both the bed and freeboard regions.Continuous addition of fresh solid particles may be required for some operations(e.g., coal gasification).

The performance is typically described using both a fluidized bed model and a freeboard-reaction model. Applications of fast-fluidized beds are in fluidized-bed combustion and Fischer-Tropsch synthesis of hydrocarbons from CO and H2.

In pneumatic flow, fluid velocities are considerably greater than the terminal velocities of the particles, so that virtually all of the particles are entrained. The vessel may be extremely tall, with no solid recirculation (e.g., coal combustion), or it may provide for solid recirculation with external cyclones.

The process stream is extremely dilute in solid particles because of the high volumeof gas passing through the “bed.” Fluid-catalyzed cracking of gasoil is an important example of pneumatic transport with external recirculation (and regeneration) of catalyst pellets.

A spouted bed is characterized by a high-velocity spout of gas moving up the center of the bed, carrying particles to the top. This action induces particle circulation, with particle motion toward the wall and downward around the spout and toward the center. The particles in a spouted bed are relatively large and uniformly sized.

Spouted beds have been used for drying and low-temperature chemical-treatment. Examples are the low-temperature roasting of agricultural products, and particle-coating and crystal-growth operations.

EXAMPLES OF REACTIONS Reactions in moving-particle reactors in general, and in fluidized-bed reactors in

particular, may be catalytic or noncatalytic. That is, the particles may be catalyst particles or reactant particles.

Catalytic cracking of gasoil: an impetus for the development of fluidized-bed reactorsover 50 years ago was the desire to make the catalytic cracking of gasoil(to gasoline) a continuous process, in spite of the rapid deactivation of the catalyst particles by coke and tarry deposits. Originally, both the catalytic-cracking reactor (“cracker”) itself and the catalyst regenerator were fluidized-bed reactors,with solid particles moving continuously between the two in an overall continuous process, but more recently the cracker is made a pneumatic-transport reactor.

Production of acrylonitrile by ammoxidation of propylene (SOHIO process): 2NH3+ 3O2 + 2C2H2 -> 2C2H2N + 6H2O.The fluidized-bed process for this reaction has several advantages over a fixedbed process. First, the process is highly exothermic, and the selectivity to C,H,N is temperature dependent. The improved temperature control of the fluidized bed operation enhances the selectivity to acrylonitrile, and substantially extends the life of the catalyst, which readily sinters at temperatures in excess of 800 K.

Oxidation of napthalene to produce phthalic anhydride: C10H10 + 02 à C10H10O3 à CO + H2 Proper control of temperature is required to limit napthaquinone production and avoid the runaway (and possibly explosive) reaction which leads to theproduction of CO, and H,O. A fluidized-bed is thus preferred over a fixed-bed process.

Production of synthetic gasoline by the Fischer-Tropsch process:nCO + 2nH2 à (C,H,) + nH2O

This is another example of a highly exothermic process which requires strict temperature control to ensure appropriate selectivity to gasoline, while limiting the

production of lighter hydrocarbons. Again, the enhanced temperature control provided by a fluidized-bed system greatly improves the feasibility of this process.

Noncatalytic roasting of ores such as zinc and copper concentrates:2ZnS + 302 à 2ZnO + 2S02

The fluidized-bed process replaced rotary kilns and hearths; its primary advantages are its higher capacity and its lower air requirement, which leads to a product gas richer in SO, for use in a sulfuric acid plant.

Noncatalytic complete or partial combustion of coal or coke in fluidized-bed combustors:

C + O2 à CO2 C+1/2O2 àCO

These reactions may serve as a means of regeneration of coked catalysts. Both reactions are exothermic, and the improved temperature control provided by a fluidized

bed is critical for regeneration of catalysts prone to sintering.

The advantages and disadvantages of moving-particle reactors may be considered relativeto the characteristics and operating conditions of fixed-bed reactor.

Advantages(i) Mode of operation: operation can be made continuous with respect to both the processing fluid and the solid; this allows, for example, for the continuous regeneration of a deactivating catalyst.(ii) Thermal: there is near-uniformity of T throughout the bed, which allows for better control of T and avoidance of hot spots in highly exothermic reactions;the uniformity of T is due to such things as the high degree of turbulence(resulting in relatively high heat transfer coefficients), and the large interfacial area between fluid and small particles.(iii) Chemical performance: the use of relatively small particles (e.g., 0.1 to 0.3 mm) can result in lower pore-diffusion resistance in solid particles and an effectiveness factor (q) much closer to 1; by itself, this, in turn, results in a smaller catalyst holdup.Disadvantages(i) Mechanical: abrasion causes erosion of pipes and internal parts (e.g., heat transfer surface); attrition of particles leads to greater entrainment and elutriation,requiring equipment (cyclones) for recovery; these mechanical features lead to higher operating and maintenance costs, as well as greater complexity.(ii) Fluid-mechanical: There is a larger (-AP), requiring greater energy consumption;the complex flow and contacting patterns are difficult to treat rationally, and create difficulties of scale-up from small-diameter, shallow beds to large-diameter, deep beds.(iii) Chemical performance: in fluidized-beds, there is a “bypassing effect” which leads to inefficient contacting; fluid in large bubbles tends to avoid contact with solid particles; this leads to a larger catalyst holdup and/or lower conversion,which may even be lower than that predicted on the basis of BMF,which in turn is lower than that based on PF, the turbulence and resulting backmixing may result in adverse effects on selectivity, for example, depressing the intermediate in a series reaction.DESIGN

Typical design requirements include calculations of catalyst or reactant solid holdup for a given fractional conversion and production rate (or vice versa),

The bed depth, the vessel diameter and height, and heat transfer requirements. The reactor model may also need to account for conversion in regions of the

vessel above(“freeboard” region) and below (“distributor” region) the bed, if there is a significant fraction of the solid in these regions, and/or the reaction is very rapid.

The overall design must consider special features related to the superficial velocity, and the flow characteristics of the solid and fluid phases within the vessel.

The following regimes of (- ΔP), somewhat idealized for simplification, corresponding to different conditions of the bed may be distinguished:

AB: (- ΔP) increases with increasing u, , as the bed retains the character of a fixed bed, with l g increasing somewhat.

BC: (- ΔP) is relatively constant, as the bed becomes fluidized; B is the point of incipient fluidization at urnf.

CD: (- ΔP) decreases with increasing us, from C, at which elutriation begins, to D, at which point the particles may be said to be entrained (at ut).

Umf2 + 150 umf( 1-εmf)μf/1.75ρfdp’ – g(ρg –ρf ) εmf3dp’/1.75 ρf=0

A hydrodynamic model of fluidization attempts to account for several essential features of fluidization: mixing and distribution of solids and fluid in a so-called “emulsion region,” the formation and motion of bubbles through the bed (the “bubbleregion”), the nature of the bubbles (including their size) and how they affect particle motion/distribution, and the exchange of material between the bubbles (with little solidcontent) and the predominantly solid emulsion. Models fall into one of three classes(Yates, 1983, pp. 74-78):

(1) two-region models, which take into account a bubble region and an emulsion region, with very little variation in properties within each region;

(2) bubble models, which are based upon a mean bubble size; all system propertiesare functions of this bubble size; (3) bubble-growth models, which also endeavor to account for bubble coalescence and bubble splitting.K-L MODEL FOR BFBPass an excess of gas upwards through a bed of fine particles.As simplifications we assume the following :

The bubbles are all spherical of same size ds.The bed contains bubbles surrounded by thin clouds rising through the emulsion.

The emulsion stays at minimum fluidization conditions.Hence the relative G/S velocity stays constant.

Each bubble drags a wake of solids behind it.This generates a ciruculation of solids in the bed,upflow behind bubbles and downflow everywhere else in the bubble.This occurs when

uo > ( 3 to 11)umf uo Superficial gas velocity in the bed.

MATERIAL BALANCE FOR GAS AND SOLIDS Ubr = 0.711(gdb)1/2 Rise velocity of a single bubble ub = uo - umf + ubr Rise velocity of bubbles ∂ Bed fraction in bubbles ∂ = (uo – umf)/ub = 1 – ubr/ub Hm(1-εm) = Hmf(1- εmf) = Hf(1- εf)

1 - ∂ = (1- εf) /(1- εmf) =Hmf /Hf Us = α ∂ub / 1- ∂ - α ∂ Down flow of emulsion solids Ue = Umf/ εmf – Us Down flow of emulsion gas Using Davidson’s theoritical expression for bubble cloud circulation the

interchange of gas between bubble and cloud is found to be Kbc = 4.5( umf/db) + 5.85 (D1/2g1/4 / db5/4)

Between cloud and emulsion Kce = 6.77 (εmf Dubr/db3)1/2

fb = 0.001 – 0.01 Volume of solid in bubble

fc = ∂(1- εmf) [ 3umf/ εmf + α ] Volume of solid in cloud ----------- ubr - umf/ εmf

fe = ( 1- ∂)(1- εmf ) - fb -fc Volume of solids in emulsion Hbfb = W / ρs A ( 1- εmf )

FIRST ORDER REACTION A à R ra’’’ = k’’’Ca ln Cao/Ca = k’’’τ’’’ = [fbk’’’ + 1 ---------------------------------- 1/ ∂Kbc + 1 ------------------ fck’’’ + 1 ------------- 1/ ∂Kce +1/fek’’’ ]ftotxHbfb ---------- ftot x uo

lnCao /Ca = k Lfl/Ub H = W/ρp(1-εmf)A q = Fao RT/P D = ( 4q / π uo)1/2 Wcat = (ρpq( 1- εmf )ubr/kuo) ln (Cao/Ca)

As gas bubbles reach the upper surface of a fluidized bed, they burst, releasing gas and ejecting particles into the freeboard region above the bed. The solid particles thrown into the freeboard originate from the bubble wakes, and cover the entire range of particle sizes present within the bed.

Several models have been proposed to account for reaction in the freeboard. Yates and Rowe (1977) developed a simple model based upon complete mixing of particles in the freeboard, coupled with either BMF or PF of the freeboard gas.

Experiments have shown that the following design features influence the extent of particle entrainment, and, by extension, the likelihood of reaction in the freeboard region:

(1) Vertical internals do not affect entrainment in small-particle beds, but may increase entrainment in large-particle beds.(2) Horizontal louvers placed near the bed surface can significantly decrease entrainment in large-particle systems.(3) Entrainment increases markedly at high pressure.(4) Stirrers or bed internals which reduce the size of the bubbles bursting at the surface can significantly reduce entrainment (reducing the number of particles ejected into the freeboard).MEMBRANE BIOREACTORS

They are mainly a combination of a membrane process like micro filtration or ultra filtration with a suspended growth bioreactor

They are mainly used in industrial wastewater treatments There are two types of configurations Internal(The membrane is inside the reactor and is an integral part of the reactor) External(The membrane is a separate unit)

In the internal type of membrane bioreactors the membrane used is a hollow fibre

The hollow fibre are of different types. Asymmetric Symmetric Flat sheet

One of the reasons that hollow fibre is used is to mix liquor and to treat the effluent obtained from liquor

Membrane reactors combine a catalyst filled reaction chamber with a membrane to add reactants or remove products of the reaction

One of the important factor that has to be considered in a membrane bioreactor is

MEMBRANE FOULING It affects the system performance This leads to a increase in the transmembrane pressure Membrane cleaning is hence required,this leads to the increase in the

operating costs In the case of a submerged reactor(Internal) air induced cross flow can

efficiently remove or atleast reduce the fouling layer on membrane surface

Aeration plays a major role in reducing membrane fouling Membranes with high pore size may foul rapidly due to clogging The membrane pore size should be about0.02 to0.5mum Two main ways to reduce fouling is

Membrane Backwashing Air backwashing

Intensive cleaning can also be carried out The frequently used cleaning agents are sodium hypochlorite and

citric acid Nutrient removal can also be done by nitrification and

denitrification processes combinedAPPLICATIONS OF MEMBRANE BIOREACTORS:

Waste water treatments To remove pathogens like (cryptspordium) Produce high quality of effluent Water reclamation Mico pollutant removal ADVANTAGES AND DISADVANTAGESINTERNAL(ADVANTAGES)

Lower flux Less frequent cleaning required

Lower operating cost Very low liquid pumping costDISADVANTAGES

Aeration cost is high High capital costEXTERNAL(ADVANTAGES)

Aeration cost is low Low capital costDISADVANTAGES

High pumping cost High flux More frequent cleaning required High operating cost

Derive the Design equation for a packed bed reactor.(Also give the design considerations).

In addition to flow, thermal, and bed arrangements, an important design considerationis the amount of catalyst required (W), and its possible distribution over two or morestages. This is a measure of the size of the reactor. The depth (L) and diameter (D)of each stage must also be determined. In addition to the usual tools provided by kinetics,and material and energy balances, we must take into account matters peculiarto individual particles, collections of particles, and fluid-particle interactions, as well as any matters peculiar to the nature of the reaction, such as reversibility. Process design aspects of catalytic reactors are described by Lywood (1996).

Characteristics of a catalyst particle include its chemical composition, which primarily determines its catalytic activity, and its physical properties, such as size, shape, density,and porosity or voidage, which determine its diffusion characteristics. We do not consider in this book the design of catalyst particles as such, but we need to know these characteristics to establish rate of reaction at the surface and particle levels (corresponding to levels (1) and (2) in Section 1.3). This is treated in Section 8.5 for catalyst particles. Equations 8.5-1 to -3 relate particle density pp and intraparticle voidage or porosity p.The shape of a particle may be one of many that can be formed by extrusion or tabletting. As in Chapter 8, we restrict attention to three shapes: (solid) cylinder, sphere, and flat plate. The size for use in a packed bed is relatively small: usually about a few mm.At the bed level (corresponding to level (3) in Section 1.3) we also use characteristics of density and voidage. The volume of the bed for a cylindrical vessel is

V = πD2L/4

The bulk density, ρb, of the bed is the ratio of the total mass W to the total volume V:

ρb = W/V

The bed voidage, s, is the ratio of the interparticle void space to the total volume V:

εb =(V - volume of particles )/V= (V - V ρb/ ρp),/V = 1 - ρb /ρp

where ρp is the density of the particle, equation 8.5-1. From equations 21.3-2 and 8.5-3,

ρb= ρp(1- εb) = ρs(1- εb) (1- εp)

where is the true density of the solid (no intraparticle voidage).

When a fluid flows through a bed of particles, interactions between fluid and particleslead to a frictional pressure drop, (- AP). Calculation of (- AP) enables determinationof both L and D, for a given W (or V). This calculation is done by means of the momentum balance, which results in the pressure gradient given by

dP/dx + fu2ρf /dp’ = 0

where x is the bed-depth coordinate in the direction of flow, f is the friction factor, u is the superficial linear velocity, pf is the density of the fluid, and d; is an effective particle diameter. Integration of equation 21.3-4 on the assumption that the second term on the left is constant results in the pressure drop equation: -

-ΔP/L = fu2ρf /dp’

where L is the depth of the bed (L = x at the outlet); note that ΔP = P(outlet) -P(inlet) is negative. In equation 21.3-5,dp; is the effective particle diameter, which accountsfor shape, defined asdp = 6 X volume of particle/external surface area of particleFor a spherical particle of diameter dP, dp’ = dP, and for a solid cylindrical particle,dh = 3d,l(2 + d,lL,) or 1.5d,, if dp/Lp << 2, where L, is the length of the particle.For the friction factor, we use the correlation of Ergun

f = [1.75 + 150(1-εp) /Re’ ] ( 1- εb)/ εb3 )

where εb is the bed voidage, and Re’ is a Reynolds number defined by

Re’ = dp’uρf / μf

Re’ = dp’G / μf

where μf is the viscosity of the fluid, and G is the mass velocity:

G = 4 m /πD2

where m is the mass flow rate, and A, is the cross-sectional area of the bed.If D or L is specified (for a given W), (-ΔP) may be calculated from above 5 equations. Alternatively, we may use these equations to determine D (and L) for a specifiedallowable value of (-ΔP ). The equations may be solved explicitly for D (or L):

D6 –βКD2 – К = 0

β = 150 π μf( 1- εb)/4 x1.75 dpm =67.32 μf( 1- εb)/dpm

К = 64 αm2W/ π3 ρf dp(-ΔP) ρb

α = 1.75( 1- εb)/ εb3

Above Equation provides the value of the bed diameter D for a given allowable pressuredrop, (-ΔP), a value of W calculated as described later, and knownvalues of the other quantities. Since (α,β, and К are all positive, from the Descartes ruleof signs there is only one positive real root of the above equation . If theequations are solved for L instead of D, a cubic equation results.In the choice of a value for the allowable (-ΔP) on the one hand, or for D (or L)on the other (given W), there is a trade-off between the cost of the vessel and the costof pumping or compressing the fluid. The smaller D, the greater the L/D ratio and thegreater (-ΔP); thus, the cost of the vessel is less, but the cost of pumping is greater, andconversely.

The process design of an FBCR involves exploiting the continuity (material-balance)and energy equations to determine, among other things, the amount of a specifiedcatalyst required for a given feed composition, fractional conversion, and throughput;concentration and temperature profiles; and the thermal mode of operationto achieve the objectives. The appropriate forms of these equations, together withrate equations for reaction and heat transfer, constitute the main working equationsof a reactor model. Before using any particular model for calculations, wedescribe a classification as a basis for consideration of models in general for anFBCR.The basis for a classification is shown in Figure below, and control volumes are shownin Figure below for axial flow. These diagrams could apply to catalyst packed inside avessel or inside a tube in a multitubular arrangement

PACKED BED REACTORSGas can be made to contact the solid in many ways giving many contacting patterns. These may be divided into two categories.

Fixed Bed Reactors

Fluidized Bed Reactors

The moving bed reactor is the intermediate case which embodies the merits and demerits of fixed and fluidized bed reactors.

MERITS AND DEMERITS In passing through the fixed beds gases approximate to plug flow whereas the

flow is complicated for fluidized beds. This behavior leads to ineffective contacting and requirement of more catalyst for high conversion. For effective contacting fixed bed reactor is favored.

Effective temperature control of large fixed beds is difficult. In contrast rapid mixing of solids in a fluidized bed allow isothermal operations. If operations are restricted to a narrow temperature range, fluidized bed reactor is preferred.

Fixed beds cannot use very small size particles because of plugging whereas fluidized bed can use small size particles.

If catalyst is to be regenerated frequently then fluidized bed allows it to be pumped from unit to unit.This feature of fluidzed bed offers overwhelming advantage over fixed bed.

DESIGNMany factors must be considered for optimum design and the best design is the one which uses two different reactor typesin series.

The main difficulties in design of catalytic reactors is

How to account for the non isothermal behavior for packed beds. How to account for non ideal flow of gas in the fluidized beds.

STAGED ADIABATIC PACKED BED REACTORS With proper interchange of heat and proper gas flow,staged adiabatic packed bed

systems become a versatile system which is able to approximate any desired temperature progression.

The reacting conditions should follow optimal temperature progression. For any preset number of stages the optimizing operations reduces to minimizing

the total catalyst for a given conversion.

OPTIMUM TWO STAGE PACKED BED REACTOR

In searching for the optimum we have three variables to set ; the incoming temperature ( point Ta),the amount of catalyst used in the first stage(locates point b along adiabatic),the amount of intercooling (locates point c along bc line).

We are able to reduce this 3-d search to 1-d search where Ta alone is guessed.

The procedure is as follows.

Guess Ta Move along the adiabatic line till the following condition is satisfied.

out ∫ ∂/∂T (1/-ra’)dXa = 0 in

Cool to point c which has the same rate of reaction as point b thus, (-ra’ )leaving the reactor = ( -ra’)entering next reactor

Move along the adiabatic from point c till the criterion given in second point is satisfied giving point d.

If point d is at desired conversion then Ta guessed is correct

Else guess another value of Ta and repeat the above steps.

For three or more stages just extend the above procedure.CHOICE OF CONTACTING SYSTEM

For endothermic reaction rate always decreases with conversion and hence we should always use plug flow with no recycle.

All else being equal cold shot cooling has the advantage of lower cost since heat exchangers are not required.Cold shot cooling is practical only when feed temperature is very much below the reaction temperature.

For exothermic reactions if the slope of adiabatic line is low it is advantageous to avoid low temperature regime.Thus high recycle approaching mixed flow is favored.If slope is high plug flow should be used.

For exothermic reactions For pure gas use high recycle rate. For dilute gas requiring no large preheating use plug flow. For dilute gas requiring large preheating use cold shot operation.

Packed (fixed) bed reactors:Packed towers are tubular reactors which are filled with packing and are used for continuous contact of liquid and gas in both countercurrent and co-current flow. The purpose of a packed bed is typically to improve contact between two phases in a chemical process. Packed beds can be used in a chemical reactor (distillation process) or in bioreactors where immobilized enzymes are used.

Packings are of two types:1. Random/unstructured – Random packings are simply dumped into the tower

during installation and allowed to fall at random. They are inexpensive and offer large surface area but lead to high pressure drops. Example: raschig ring, berl saddle.

2. Regular/structured- They are arranged in a uniform manner. They are expensive but offer the advantage of low pressure drop. Example: raschig rings stacked uniformly, metal sheets or gauzes.

Packing supports: The support must be sufficiently strong to carry the weight of a reasonable height of packing and it must have ample free area to allow for flow of liquid and gas with a minimum of restriction.Packing restrainers: These are necessary when gas velocities are high and they guard against the lifting of packing during a sudden gas purge. Heavy screens or bars may be used.

Immobilized enzymes in packed bed reactors(PBR):

Enzyme (biocatalyst) immobilization is done so that their activity can be retained and they can be reused in the bioreactors. Various methods such as gel entrapment, cross linking, covalent binding etc are used for immobilization. Like other bio reactors immobilized enzymes are used in packed bed reactors for processes like glucose isomerization, selective penicillin hydrolysis, and for selective reactive separation of racemic mixture of amino acids. Generally enzymes are immobilized in calcium alginate and then packed into the bioreactors. In the PBR, the feed (fluid) is allowed to flow through these packings consisting of the immobilized biocatalysts. Metabolites and products are released into the fluid and removed in the outflow. The feed velocity (superficial velocity) is below the minimum fluidization velocity.

The flow might be upwards or downwards, but upflow is more desirable as the fluid will contact/wet the packing uniformly (avoids short circuiting/channeling) and also the contact time will be more than that of downflow.The PBR follows the plugged flow pattern, thereby only radial mixing will occur and there won’t be any axial mixing (/backmixing) inside the reactor. Hence the residence time distribution will be uniform as all the fluid particles literally travel the same path length inside the reactor.The substrate concentration will be more near the fluid inlet and the product concentration will be less. As the reaction progresses the product concentration increases and hence near the outlet the substrate concentration will be less and the product concentration will be more. This important characteristic makes the PBR useful for carrying out product inhibition reactions.In product inhibition reactions, there will be competitive inhibition as the product present initially competes with the substrate for binding with the enzyme. However, as only a small portion of product will initially compete with the substate in the PBR, the product inhibition reactions can successfully be carried out in these reactors.Pressure drop:It is the decrease in pressure from one point to another point along the length of the reactor.Tube convergence, divergence, turns and the flow rates affect the pressure drop.It is given by the Ergun equation:

High flow rates in small tubes give larger pressure drop, low flow rates in large tubes give lower pressure drops. High pressure drops are undesirable because it will result in both high pumping energy costs and non-uniform pressure of the gas reactant, affecting the overall reactor performance.The pressure drop in the reactor can be reduced by using larger catalyst particles. However smaller particles will increase the surface area of contact. So the diameter of the particle used should be optimum favouring a low pressure drop as well as giving a good surface area for contact.

Performance equations:These are also called as design equations as they are useful for reactor design.Considering mass balance:Rate of substrate diffusion out of bulk liquid = rate of substrate disappearance by reaction within the pellet4πR2ks(S-Ss) = (4/3) πR3η (Ss, Ps) v (Ss, Ps)Where, R= radius of pellet (biocatalyst)ks = saturation constantS = substrate concentration in bulk liquidSs = substrate concentration on the exterior surface of the pelletη = effectiveness factorv = intrinsic rate of product formationPs = product concentration on surface of pelletConsidering material balance on substrate:u.ds/dz = - ((1-ε)/ε)η (Ss,Ps)v (Ss,Ps)Where, ε = void fraction1-ε = fraction occupied by the solid particles in the packingIf η = 1, then Ss = S; i.e. no mass transfer resistance will be there between the liquid and solid phase when effectiveness is 1.

In design, the number of necessary theoretical equilibrium stages is first determined and then the packing height equivalent to a theoretical equilibrium stage, known as ‘Height equivalent to a theoretical plate (HETP)’ is determined.Total packing height = (number of theoretical stages) x (HETP)

Advantages:1. High conversion per unit mass of catalyst

2. Low pressure drop

3. For processing product inhibition reactions

Disadvantages:1. Poor temperature control

2. Poor pH control

3. Difficult to service and clean the packings.

Residence Time Distribution for Chemical Reactors

General CharacteristicsThe two major uses of the residence time distribution to characterize nonideal reactors are1. To diagnose problems of reactors in operation2. To predict conversion or effluent concentrations in existing/available reactors when a new reaction is used in the reactor.

Three concepts were used to describe nonideal reactors: the distribution of residence times in the system, the quality of mixing and the model used to describe the system.

~The time the atoms have spent in the reactor is called the residence time of the atoms in the reactor.~In any reactor, the distribution of residence times can significantly affect its performance.~The residence-time distribution (RTD) of a reactor is a characteristic of the mixing that occurs in the chemical reactor.~Not all RTDs are unique to a particular reactor type; markedly different reactors can display identical RTDs.~The RTD exhibited by a given reactor type yields distinctive clues to the type of mixing occurring within it and is one of the most informative characterizations of the reactor.

Measurement of the RTDThe RTD is determined experimentally by injecting an inert chemical,

molecule, or atom, called a tracer, into the reactor at some time t=0 and then

measuring the tracer concentration, C, in the effluent stream as a function of time.~In addition to being a nonreactive species that is easily detectable, the tracer

should have physical properties similar to those of the reacting mixture and be

completely soluble in the mixture.~It also should not adsorb on the walls or other surfaces in the reactor. The

latter requirements are needed so that the tracer’s behavior will honestly reflect

that of the material flowing through the reactor.~Colored and radioactive materials along with inert gases are the most

common types of tracers. Pulse Input ExperimentIn a pulse input, an amount of tracer N0 is suddenly injected in one shot into

the feedstream entering the reactor in as short a time as possible. We shall analyze the injection of a tracer pulse for a single-input and single-

output system in which only flow carries the tracer material across system boundaries

First, we choose an increment of time Dt sufficiently small that the concentration of tracer, C(t), exiting between time t and t+Dt is essentially the same. The amount of tracer material, DN, leaving the reactor between

time t and t+Dt is then

For pulse injection we define

The quantity E(t)dt is the fraction of fluid exiting the reactor that has spent between time t and t+dt inside the reactor.

SLURRY REACTORS

SLURRY REACTORS The design of gas liquid solid reactor is very much dependent on the size of the

solid particles chosen for the reaction.

Smaller the size of the particle the closer is the value of effectiveness factor to 1.

Particles smaller than 1 mm cannot be used in fixed beds.

Small particles should only be used as suspensions in liquid.TYPES OF REACTOR

Three phase reactors can be mainly divided into 2 classes.

Three Phase reactor

Suspended Bed Fixed Bed Reactors Reactors

Bubble Agitated

Columns TanksBUBBLE COLUMNS AND AGITATED TANKS

If particles are very small they can be quite easily kept in suspension by bubbling gas through liquid in a bubble column.

Agitated tanks can be used for large size particles because of the high local velocities produced by the impeller.

The impeller needs to be properly designed and positioned at the bottom of the tank so that it will keep the solids in

suspension and disperse the incoming gas. A slurry reactor is a multiphase flow reactor in which the reactant gas is bubbled

through a solution containing solid catalyst particles.

The solution may be a reactant as in the case of methyl linoleate or an inert as in the case of Fischer Tropsch synthesis of methane.

In slurry reactor the catalyst is suspended in liquid and gas is bubbled through the liquid.

The slurry reactor may be operated either in semibatch or continuous mode.

In addition constant overall catalytic activity can be maintained by addition of small amount of catalyst with each reuse in the batch operation or constant feeding in continuous opearation.

In modelling slurry reactors we assume that the liquid phase is well mixed and the catalyst particles are uniformly distributed and gas phase is in plug flow

SLURRY REACTOR - DIAGRAM

RATE EQUATION

Let the gas consist of pure reactant A ( typically H2) and let the reaction take place at the interior of the catalyst, the reaction being pseudo first order with rate constant k1.

In an agitated tank suspended bed reactor gas is dispersed as bubbles and it is assumed that liquid phase is well mixed and concentration of A in liquid CAL is uniform throughout.

In this case the overall reaction rate Rt will be based on the unit volume of the whole dispersion ( Gas + Liquid +Solid )

Let a : Gas liquid interfacial area per unit volume of the dispersion. εg = volume fraction of gas bubbles. εl = volume fraction of liquid. εs = volume fraction of solid.

εg + εl + εs = 1

For particles external surface is given by ap = External surface area per unit volume of particle

The above quantities can be related to volume/surface or sauter mean bubble / particle diameters by the following relations.

db = 6 εg /a for bubbles dp = 6/ap for particles

The surface area of the particles per unit volume is apεp

The above quantities can be related to volume/surface or sauter mean bubble / particle diameters by the following relations.

db = 6 εg /a for bubbles dp = 6/ap for particles

The surface area of the particles per unit volume is apεp

The reactants in the gas phase participate in 5 reaction steps.

Absorption from gas phase into liquid phase at bubble surface.

Diffusion in the liquid phase from bubble surface to bulk liquid.

Diffusion from bulk liquid to external surface of solid catalyst.

Internal diffusion of reactants into the porous catalyst.

Reaction within the porous catalyst. Each step may be thought as a resistance to the overall rate

of the reaction.

The concentration in the liquid phase is related to gas phase concentration by Henry’s law. Ci = Pi H

Consider hydrogenation of methyl linoleate to form methyl oleate. Methyl Linoleate + Hydrogen à Methyl Oleate (L) (H2) (O)

Hydrogen is absorbed in liquid methyl linoleate diffuses to the external surface of the catalyst pellet,then diffuses into the catalyst pellet where it reacts with methyl linoleate to form methyl oleate. Methyl oleate diffuses out of the pellet back to the liquid.

The steps 1 and 2 and 4 and 5 can be combined to single steps.

The mass transfer and chemical reaction steps take place in series. Hence each step should proceed at the same rate as the

overall process. For gas-liquid mass transfer Rt = Kla ( Cai –Cal) For liquid-solid mass transfer Rt = Ksap εp(Cal –Cas)

For rate of reaction within the particle Rt = k1Casƞ εp

Since the rate of the reaction is the same the overall rate of the reaction can be obtained as

Rt = CAi / ( 1/Kla + 1/ Ksapεp + 1/K1 ƞεp )

1/Kla - Resistance for gas-liq mass transfer step1/ Ksapεp - Resistance for liq-solid mass transfer step1/K1 ƞεp - Resistance for pore diffusion with reaction stepDESIGN OF SLURRY REACTOR

Thiele Modulus φ = (Vp /Sp )*√k/De Vp /Sp = ro / 3 For spherical particles ƞ = 3/ λro ( coth λro – 1/ λro ) λ = √k/De STEPS 1. Calculate λro and ƞ.(From given k,De and ro)

2. The size of the bubbles produced in the reactor and gas volume fraction will depend on agitation conditions and the rate at which gas is fed to impeller.

Typical values are db =0.8 mm and εg = 0.203. The mass transfer coefficients Kl and Ks depend on the physical properties of the system such as viscosity of the liquid and diffusivity of the dissolved gas. The corrleations are For no Shear : Ks = 2DA-B/dp For Shear : Ksdp/DA-B = 2 + 0.6(Re)1/2(Sc)1/3 Ks α (U/dp)1/2 Typical values : Kl =1.23 x10-3 and Ks =0.54 x10-3

4. The most important parameter to be specified is catalyst loading ie the ratio of catalyst to liquid to be charged into the Reactor.A solid/liquid ratio of about 0.1 can be used.

5. The pressure of the gas in the reactor also needs to be fixed.6. Using the equations , db = 6 εg /a for bubbles dp = 6/ap for particles we can calculate a and ap7. Using εp / εl = 0.1 and εg + εl + εp =1 we can calculate εl and εp.

8. We can calculate 1/Kla , 1/ Ksapεp and 1/K1 ƞεp and hence the overall rate of the reaction can be determined. 9. Mixed flow of G and any flow of L A material balance for A and B about the reactor as a whole yields Fao Xa= Fbo Xb/ b = (-ra’’’’ )Vr

Fao Xa - Rate of loss of A. Fbo Xb/ b - Rate of loss of B(-ra’’’’ )Vr - Rate of disappearance of A by reaction

The above equation is derived as follows.

Fao = Fa + (-ra’’’’ )Vr

Fao - Fa = (-ra’’’’ )VrFao – Fao ( 1-Xa) = (-ra’’’’ )VrFaoXa = (-ra’’’’ )Vr

Similarly ,FboXb = (-rb’’’’ )Vr (-ra’’’’ ) = (-rb’’’’ )/b

FaoXa = FboXb/b = (-ra’’’’ )Vr

From the above equations the volume of the reactor can be foundout knowing the value of the rate of the reaction.10. Mixed flow of G and batch L. For component A Input = Output + Disappearance Fao = Fao(1- Xaexit) + (-ra’’’’)Vl For component B Input = Output + Disappearance + Accumulation 0 = 0 + dNb/dt + (-rb)Vr (-rb)Vl = -dNb/dt -rb = (1/ Vl ) -dNb/dt = -dCb/dt

- ra = -rb/ b -rb = b(-ra) = -dCb/dt -ra = 1/ b (-dCb/dt) (-ra)Vl = (-ra’’’’)Vr Thus Vl/b (-dCb/dt) = (-ra’’’’)Vr = FaoXaexitUSES

Hydrogenation of 1. Fatty acids over supported nickel catalyst. 2. 2-butyne-1,4-diol over pd-CaCO3 catalyst.

Oxidation of 1. C2H4 in inert liquid over a pdCl2 –Carbon catalyst. 2. SO2 in inert water over activated carbon catalyst.

Hydroformation 1.CO with high molecular weight olefins with cobalt complex bound polymers.

Ethynylation 1. Reaction of acetylene with formaldehyde over a CaCl2 supported catalyst. TANK REACTORS

Tank reactors usually employ mechanical agitation to bring about more intimate contact of the phases, with one phase being dispersed in the other as the continuous phase.

The gas phase may be introduced through a “sparger” located at the bottom of the tank;this is a circular ring of closed-end pipe provided with a number of holes along its length allowing multiple entry points for the gas.

Tank reactors are well suited for a reaction requiring a large liquid holdup or a long liquid-phase residence time. The operation may be continuous with respect to both phases, or it may be semicontinuous (batch with respect to the liquid).

The simplest flow pattern for each phase is BMF. The reactor may be single-stage or multistage and the flow for the latter may be cocurrent or countercurrent.

In the case of a liquid-liquid reaction, each stage typically consists of a mixer (with agitation, for reaction) and a settler (without agitation, for separation by gravity).

Tank reactors equipped with agitators (stirrers, impellers, turbines, etc.) are used extensively for gas-liquid reactions, both in the traditional chemical process industries and in biotechnology.

In comparison with nonagitated tank reactors equipped only with spargers, mechanically agitated tank reactors have the advantage of providing a greater interfacial area for more efficient mass transfer

The enhanced mixing also ensures a nearly uniform temperature within the vessel, an advantage for processing temperature-sensitive materials, and for control of product yield and selectivity in complex systems.

A major disadvantage is the cost of the energy required for agitation.

Furthermore, since the mass transfer characteristics are related to agitation characteristics (e.g., stirrer rate), there are difficulties of scale-up.

CHOICE OF TOWER OR TANK REACTOR The choice between a tower-type and a tank type reactor for a fluid-fluid reaction

is determined in part by the kinetics of the reaction. As described by the two-film model for gas-liquid reactions the rate of the overall

process is influenced by relative rates of mass transfer and of intrinsic reaction. The two extremes , for a nonvolatile liquid-phase reactant, are virtually

instantaneous reaction in the liquid-film, which is controlled by interphase mass transfer and “very slow” reaction, which is controlled by reaction itself in the bulk liquid

Typical values of gas-liquid interfacial area (ai and ai’) for various types of vessels

Tower(Spray,Plate) Tank Agitated Sparger ai 1000 200 20ai’ 100 180 20vl/vr 0.1 0.9 0.98

- ai interfacial area based on unit volume of liquid phase, m2 m-3 (liquid)- ai’ interfacial area based on unit volume of vessel (occupied by fluids),m2 m-3 (vessel)

The two quantities ai and ai’ are related by ai’ =(Vl/Vr)ai, = (1 – εg)ai where Vl, = volume of liquid in the vessel, Vr, = volume of reactor (vessel) occupied by fluid, and εg is the gas holdup, m3 gas (m3 reactor)-1.

Values of the ratio Vl / Vr given in table emphasize that most of the volume in a tower reactor is occupied by the gas phase, and conversely for a tank reactor. This means that ai >> ai’ in a tower and ai= ai’ : in a tank.

For a tank reactor the ideal flow pattern is BMF for each phase (gas and liquid). Important considerations are the dispersion of gas bubbles within the liquid phase, and the agitator power required for this.

The h/D ratio is usually about 1, but this is exceeded if the stages in a multistage arrangement are stacked vertically. Gas is usually fed to the reactor below the agitator through a distributor which may be a perforated plate or pipe ring (sparger).

The type of agitator also can affect bubble size, and, thus, mass transfer characteristics.

The most common types are rotating flat paddle mixers, with straight or inclined blades, and a turbine, with several flat blades attached to a disc (turbine).

The diameter of the blades is usually about 35% of the vessel diameter (D). Several vertical baffles, each with a width of up to 10% of D, are attached to the wall of the vessel around its circumference.

CONTINUITY EQUATION Continuity equation for A in the gas phase (BMF).Since the gas phase is in BMF,

the continuity equation corresponding to and based on the entire vessel of volume V=Ach = πd2/4 h

Rate of input = Rate of output + Rate of transfer of A by bulk flow of A by bulk flow transfer of A to liquid film YainGAc = YaoutGAc + Naai’Ach

Continuity equation for A in liquid phaseRate of input + Rate of transfer of A = Rate of output of A by bulk flow from liquid film of A by bulk flow + Rate of reaction of A in bulk liquid

Cainql +ai’NaV = Caoutql + (1-εg)(-raint)V

Overall material balance around tank

The overall material balance around the tank is again given by equation G/P ( pain –paout ) = L/b ( Cbin –Cbout) + L(Caout-Cain)

CORRELATIONS FOR DESIGN PARAMETERS FOR TANK REACTORS Before we can apply the continuity or design equations for tank reactors

developed we must have the means of determining values of the parameters involved. These include gas holdup, g, mass transfer coefficients, kAl , and kAs, and gas-liquid interfacial area, ai or ai’.

For a nonelectrolyte liquid phase, the correlation of Robinson. (1977) is εg = 0.21( qgnimp2/σ)0.57

Power Input Pl = 0.34nb,imp1/2 ( Plo nimpDimp3/qg)m

Plo Power Input without gas flowlnb,imp Number of impeller bladesNimp Rate of rotation of the impellerDimp Diameter of impeller m =0.45 for non coalescing liquids =0.33 for ionic liquids

Chandrasekharan and Calderbank (1981) proposed the following correlation, which shows a much stronger inverse dependence on vessel diameter:

kalai’ = 0.048 (Pl/vl)0.5 qg0.55D-0.5

The correlation of Calderbank (1958) for ai’ isai’ = 22.8 ( Pl/Vl)0.4(usg/ubr) (ρl/σ)3

L(D) = ql /Ac

G(D) = Fain,/0.21Ac

V(D) = (1 - εg)V

Usg(D) = qg/Ac

UNIT III BIOREACTOR SCALE – UP 9Regime analysis of bioreactor processes, oxygen mass transfer in bioreactors -microbial oxygen demands; methods for the determination of mass transfer coefficients;mass transfer correlations. Scale up criteria for bioreactors based on oxygen transfer,power consumption and impeller tip speed.

Oxygen Transfer in Bioreactors

Oxygen is needed by cells for respiration. Oxygen used by cells in suspension must be available as dissolved oxygen. Since oxygen solubility is quite small, about 6 to 7 mg/L under normal cultivation conditions, metabolic oxygen requirement is supplied on a as needed basis by continuous aeration of culture medium. Actively respiring yeast requires about 0.15 g O2 (g cell)-1 h. At a cell concentration of 10 g L-1, medium saturated with air can support less than 30 seconds worth of metabolic oxygen. That is, a continuous supply of oxygen must be maintained in any viable aerobic manufacturing process. In this Chapter, we will first get a quantitative appreciation for metabolic oxygen demand, followed by methods used in calculating rates at which oxygen is transfered from sparged air. We will then examine methods useful in characterizing oxygen mass transfer coefficient. Finally we will evaluate bioreactor operation and design based on oxygen transfer capability.

 Metabolic Oxygen Demand

Metabolic oxygen demand of an organism depends on the biochemical nature of the cell and cultivation conditions. Oxygen need is usually satisfied in most cells if the dissolved oxygen concentraiton in the medium is kept at about 1 mg/L. If the oxygen level is allowed to fall far below this value, oxygen consumption rate decreases with concomitant decrease in biochemical energy production, and as a result cell growth rate also decreases. We described this behavior in Section 4-4. The value of oxygen concentration above which growth rate is at the maximum was described as the critical oxygen

concentration, . Characteristic values are summarized in Table 5-1.

Table 5-1 Critical Oxygen Concentration

Organism in mg L-1

E. coli at 37 C 0.26

S. cerevisiae at 30 C 0.13

Penicillium sp at 24 C 0.78

The oxygen requirement for growth is expressed best in the the parameter, yield coefficient, YX/O2. It represents the amount of oxygen required to grow one gram of cells. Typical values summarized in Table 5-2, show that approximately 0.7 to 1 g of oxygen is needed to produce 1 g of cells. In the same table respiration quotient is also included.

Table 5-2 Stoichiometric Oxygen Demand &Respiration Rate

Organism Substrate YX/O2

[g (g cell)-1 h] qO2

[g O2 (g cell)-1 h]

E. coli Glucose 1.1 0.20

S.cerevisae Glucose 0.98 0.30

Candida utilis Glucose 1.32

Pennicillium sp. Glucose 1.35 0.18

Hybridoma

CHO cell line

Volumetric Oxygen Mass Transfer Coefficient

In a typical aeration system, oxygen from the air bubble is transferred through the gas-liquid interface followed by liquid phase diffusion/bulk transport to the cells. Although this is a multi-step serial transport, in a well dispersed systems, the major resistance to oxygen transfer is in the liquid film surrounding the gas bubble. Consider the oxygen concentration profiles in the region near the interface illustrated in Figure 5-1.

Figure 5-1 Oxygen Concentration Profile at Air Bubble-Medium Interface

The transport of oxygen through the gas and liquid films are equal at steady state. They can be expressed by

where subscript G and L refer to gas and liquid phases respectively. The terms, NO2G and NO2L are oxygen transfer expressed in g O2 h-1, A is interfacial area and CDO is oxygen concentration expressed in g O2 per unit volume. At the interface, equilibrium between the liquid and gas phase oxygen is reached. That is

Because of low oxygen solubility and the fact that kG is much higher than kL,

Hence, Eq (5-1a) can be written as

The subscript L in NO2 has been dropped to note that the above represents overall transfer of oxygen. The driving force in the above consists of the difference between bulk oxygen concentrations in the two phases; the first term represents the concentration of oxygen in the liquid which is in equilibrium with the bulk gas phase oxygen. If air is the gas medium, this term will equal to 7 mg/L at 35 C.

When the above oxygen transfer is applied to an entire volume of a bioreactor, A will represent the total interfacial area and kL will represent an average mass transfer coefficient. The concentrations will be bulk gas and liquid phase oxygen concentrations. If we divide the above equation by volume of liquid phase, V, the resulting term will represent the amount of oxygen transfered per unit volume per unit time --- which is in the same units as the rate expressions we saw in last chapter. Since the rate is due to a physical phenomena, let us distinguish it by the symbol, RO2. That is,

The term, kL A represents the product of mass transfer coefficient and interfacial area available for mass treansfer. In a bioreactor, air is sparged and the liquid is agitated to break up the bubbles so that interfacial area can be kept high to enhance rate of oxygen transfer. In such systems, the area, A, is not easily measured or estimated. But, the term consisting of the product - mass transfer coefficient and interfacial area - is more readily measured. Further more, it is convenient to use interfacial area per unit volume, a, rather

than total area, A because rate of oxygen transfer is expressed per unit volume of bioreactor, similar to rate of cell growth, which is reported on a volumetric basis. Hence, area per unit volume, a, is combined with the mass transfer coefficient, kL and is given by

the term, kLa. In Eq(5-5) the term, can be replaced by oxygen solubility at bioreactor conditions, .

The above will be our working equation for describing transfer of oxygen from gas phase to growth medium. In order for us to calculate oxygen transfer rate (OTR), we need the mass transfer coefficient, kLa , solubility of oxygen in the medium, and the dissolved oxygen concentration in the medium, CDOL. In the last chapter we had used the notation, CDO to describe dissolved oxygen concentration. In the discussion above, there was a need to make a distinction between gas and liquid phase concentration. In Eq (5-5), one notes that both concentrations are expressed on the basis of liquid phase. Hence, from here on we will drop the subscript L. In situations where we need to make a distinction between the two phases, we will re-introduce the subscript L and G.

Bioreactor Oxygen Balance

Let us now consider the case of oxygen balance within a bioreactor in which cells are growing and in the process consuming oxygen. There is a continuous inflow of air at a constant volumetric flow rate. The liquid broth is agitated by a Rushton agitator (flat blade stirrer ). Le the metabolic oxygen uptake rate be qO2 and cell concentration is X. Let us examine the reactor system over a sufficiently short period that we can treat X as a constant. Consider oxygen balance over the liquid phase of the bioreactor.

O2 transfered from Gas Phase - O2 consumed by Cells = Accumulation

For constant liquid phase volume, the above can be simplified to

The concentration, CDO is readily measured using an dissolved oxygen electrode. A later segment of the course on Biosensors, will deal with principle of measurement and construction of DO electrodes.

If oxygen being supplied is in exact balance with the oxygen consumed by the cells, we expect the dissolved oxygen concentration to remain constant; that is, the derivative in Eq(5-7) will vanish. That is,

One useful application of the above is in estimating the maximum cell concentration a particular bioreactor is capable of supporting in terms of oxygen supply. See the example below.

Example 5-1.

A bioreactor has an oxygen mass transfer coefficient capability of 400 h-1. What is the maximum concentration of E. coli that can be grown aerobically in this reactor. Respiration rate of E. coli is 0.35 g O2 (g Cell)-1 h-1. Critical oxygen concentration is 0.2 mg/L. Assume oxygen saturation with air to be 6.7 mg/L.

Solution

From Eq(5-8), we have

The maximum oxygen concentration driving force that can be expected is

= ( 6.7 - 0.2) = 6.5 mg/L.

Therefore, maximum cell concentration that can be grown at maximum growth rate is

Factors Affecting KLa The mass transfer coefficient is strongly affected by agitation speed and air flow rate. In general,

Note that the mass transfer coefficient increases with agitation speed and air flow rate.

Measurement of KLa

Most common method of measuring kLa is to conduct experiments in the bioreactor when cells are absent, or cell concentration is low so that consumption by cells can be neglected. The latter condition is present immediately after inoculating the bioreactor. Consider Eq (5-7) under these conditions:

If we allow steady state to occur, the dissolved oxygen concentration will reach saturation value, and the concentration-time profile will be flat, as shown in the diagram.

Fig 5-2 Oxygen Profile During a Transient. The responses will be exponential, rather than straight lines.

If the oxygen source (air) is replaced by nitrogen, the resulting response of the system is described by the above equation with the term, set to zero. That is,

The solution to the above is

If one plots the response on a semi-log plot, the slope will equal to the negative of mass transfer coefficient. It is relatively a simple experiment and the data analysis is also easy to do. When other type of transient mass transfer experiments are conducted, the above equations should be suitably modified. For example for the case of nitrogen to air switch, we should suitably modify the solution because the initial condition is now different.

Case Studies

You are part of a tech service team asked to evaluate if the available 10,000 liter fermentor is adequate to produce 10 kg/day of a recombinant protein using a strain of E. coli that expresses the protein as 20 % of cellular protein. In order to enhance plasmid stability, the nutrients are manipulated to give a low specific growth rate is 0.2 h-1. The oxygen demand is 0.15 g O2/g cell - h. Assume that the r-protein formation is cell growth associated.

Data: The lag phase is 4 hours. Typical clean-up time following a fermentation batch and preparation for the next batch is 8 hours. The plant runs three shifts. Cell yield on substrate is 0.55 g cell/g substrate. Available support services can supply inoculum of a maximum of 6 kg of cells every 24 hour period. Maximum KLa for the available

fermentor is 500 h-1. Fermentor accessories are capable of handling cell concentrations of

60 g/L. Assume any other parameters you need to complete the calculation.

Assumption: Critical oxygen conc. is 0.2 mg/L and DO at air saturation is 6.4 mg/L

Solution A: Lag phase and clean-up/ prep time is given as 12 h. If a batch is to completed within each 24 h period, production is limited to 12 h per day. If this is not a limitation, one can optimize production by varying batch time. Let us first evaluate assuming 12 h batch times.

If max. cell concentration of 20.6 g/L is obtained, amount of r-protein produced is = (0.2) (0.5) (20.6) = 2.06 g/L. 50% of cell dry matter was assumed to be protein. Hence in 10,000 liters, we will produce 20.6 kg.

Next to determine the inoculum level. The maximum batch growth phase is 12 h. Substitute in growth eqn, and assuming nutrients are present to support exponential growth during the 12 h period,

For 10,000 liters, we will need 18.7 kg every 12 h. Since only 6 kg is available, max. protein that can be produced is

{(0.2)(0.5)[0.6 Exp((0.2)(12)] 10,000 = 6.61 kg

Solution B: Now let us allow batch times to be longer than 12 h, meaning that there might not be a harvest every day. Since it is advantageous to use the max. inoculum concentration, select X0 = 0.6 g/L. This value is obtained by diving 6 kg of cells in 10,000 L. Max. cell concentration is fixed due to aeration requirements. Use the batch

growth eqn to find the batch growth time of 17.7 h. Hence 20.6 kg or r-protein will be produced every 29.7 h which gives a 24 h production rate of 16.6 kg.

What alternative way of running reactor would you recommend to achieve the production target?

You are part of a tech service team asked to evaluate if the available 10,000 liter fermentor is adequate to produce 10 kg/day of a recombinant protein using a strain of E. coli that expresses the protein as 20 % of cellular protein. In order to enhance plasmid stability, the nutrients are manipulated to give a low specific growth rate is 0.2 h-1. The oxygen demand is 0.15 g O2/g cell - h. Assume that the r-protein formation is cell growth associated.

Data: The lag phase is 4 hours. Typical clean-up time following a fermentation batch and preparation for the next batch is 8 hours. The plant runs three shifts. Cell yield on substrate is 0.55 g cell/g substrate. Available support services can supply inoculum of a maximum of 6 kg of cells every 24 hour period. Maximum KLa for the available

fermentor is 500 h-1. Fermentor accessories are capable of handling cell concentrations of

60 g/L. Assume any other parameters you need to complete the calculation.

Assumption: Critical oxygen conc. is 0.2 mg/L and DO at air saturation is 6.4 mg/L

Solution A: Lag phase and clean-up/ prep time is given as 12 h. If a batch is to completed within each 24 h period, production is limited to 12 h per day. If this is not a limitation, one can optimize production by varying batch time. Let us first evaluate assuming 12 h batch times.

If max. cell concentration of 20.6 g/L is obtained, amount of r-protein produced is = (0.2) (0.5) (20.6) = 2.06 g/L. 50% of cell dry matter was assumed to be protein. Hence in 10,000 liters, we will produce 20.6 kg.

Next to determine the inoculum level. The maximum batch growth phase is 12 h. Substitute in growth eqn, and assuming nutrients are present to support exponential growth during the 12 h period,

For 10,000 liters, we will need 18.7 kg every 12 h. Since only 6 kg is available, max. protein that can be produced is

{(0.2)(0.5)[0.6 Exp((0.2)(12)] 10,000 = 6.61 kg

Solution B: Now let us allow batch times to be longer than 12 h, meaning that there might not be a harvest every day. Since it is advantageous to use the max. inoculum concentration, select X0 = 0.6 g/L. This value is obtained by diving 6 kg of cells in 10,000 L. Max. cell concentration is fixed due to aeration requirements. Use the batch growth eqn to find the batch growth time of 17.7 h. Hence 20.6 kg or r-protein will be produced every 29.7 h which gives a 24 h production rate of 16.6 kg.

What alternative way of running reactor would you recommend to achieve the production target?

UNIT IV MODELLING AND SIMULATION OF BIOPROCESSES 12Study of structured models for analysis of various bioprocess – compartmental models,models of cellular energetics and metabolism, single cell models, plasmid replication andplasmid stability model. Dynamic simulation of batch, fed batch, steady and transientculture metabolism.

Compartmental models

If metabolite corrected arterial blood curve and dynamic PET data are available from the time of injection to the time where all important changes in tracer kinetics have been seen, and the data is of sufficient quality, then it may be possible to estimate the parameters of a complete model, including perfusion, blood volume in tissue vasculature, transport, specific binding or reaction rate etc. In practice this can be done only in very simple cases, e.g. with labeled water which just flushes in and out of tissue. As is mentioned above, it is possible and desirable to either measure some of the parameters in separate studies, e.g. blood volume using [15O]CO, or constrain them to literature values, or to reduce the model. Most often there is no need to measure all the parameters, but just one key parameter which correlates with the desired property in usual conditions.

Plasma compartment

To be precise, the plasma is not a compartment of the model. The concentration of tracer in the plasma is measured, and applied to the compartment model as a known input function. However, plasma compartment is still usually counted as one of the compartments.

The metabolite corrected arterial plasma curve is the input to the compartment model. If the intravascular activity is accounted for in the calculation, the whole blood concentration, containing metabolites, should be used for that. If metabolite corrected plasma curve is used instead of uncorrected whole blood curve to correct for vascular blood volume fraction, blood contribution at late times is underestimated, which could result in the artificial presence of an apparent additional tissue compartment [Lammertsma 2002].

Two-compartment model (one-tissue compartment model)

Methods for quantitation of perfusion (blood flow, f) with freelydiffusable tracers are based on Kety's analyses of the principles of inert gas exchange. Tracers such as [15O]H2O, [15O]butanol, [11C]butanol and [18F]fluoromethane are used for this purpose. Also a single breath inhalation of [15O]CO2 produces an arterial bolus of [15O]H2O.

Three-compartment model (two-tissue compartment model)

The kinetic model for measurement of glucose transport and phosphorylationrate in brain is based on a three-compartment model.

Four-compartment model (three-tissue compartment model)

The description of the time course of the ligand in tissue rquires a modelthat accounts for the different components contributing to the signal. These are free ligand in plasma, free ligand in tissue, CF, ligand in tissue which is not specifically bound, CNS, and ligand specifically bound to the receptor, CB [Schmidt and Turkheimer 2002]:

Use of 4-compartment model is not feasible given the large number of rate constants to be estimated, and the kinetically undistinguishable compartments for specific and non-specific binding. The model is simplified under the assumption of a rapid equilibrium between free and non-specifically bound compartments that produces a single compartment of free + non-specifically bound ligand:

Although for most ligands k3' and k4 cannot be estimated with a reasonable degree of accuracy, their ratio is usually more stable.

UNIT V BIOREACTOR CONSIDERATION IN ENZYME SYSTEMS 8Analysis of film and pore diffusion effects on kinetics of immobilized enzyme reactions;

formulation of dimensionless groups and calculation of effectiveness factors. Design ofimmobilized enzyme reactors – packed bed, fluidized bed and membrane reactors.

Enzyme reactors

An enzyme reactor consists of a vessel, or series of vessels, used to perform a desired conversion by enzymic means. A number of important types of such reactor are shown diagrammatically in Figure 5.1. There are several important factors that determine the choice of reactor for a particular process. In general, the choice depends on the cost of a predetermined productivity within the product's specifications. This must be inclusive of the costs associated with substrate(s), downstream processing, labour, depreciation, overheads and process development, in addition to the more obvious costs concerned with building and running the enzyme reactor. Other contributing factors are the form of the enzyme of choice (i.e. free or immobilised), the kinetics of the reaction and the chemical and physical properties of an immobilisation support including whether it is particulate, membranous or fibrous, and its density, compressibility, robustness, particle size and regenerability. Attention must also be paid to the scale of operation, the possible need for pH and temperature control, the supply and removal of gases and the stability of the enzyme, substrate and product. These factors will be discussed in more detail with respect to the different types of reactor.

Batch reactors generally consist of a tank containing a stirrer (stirred tank reactor, STR). The tank is normally fitted with fixed baffles that improve the stirring efficiency. A batch reactor is one in which all of the product is removed, as rapidly as is practically possible, after a fixed time. Generally this means that the enzyme and substrate molecules have identical residence times within the reactor, although in some circumstances there may be a need for further additions of enzyme and/or substrate (i.e. fed -batch operation). The operating costs of batch reactors are higher than for continuous processes due to the necessity for the reactors to be emptied and refilled both regularly and often. This procedure is not only expensive in itself but means that there are considerable periods when such reactors are not productive; it also makes uneven demands on both labour and services. STRs can be used for processes involving non-immobilised enzymes, if the consequences of these contaminating the product are not severe. Batch reactors also suffer from pronounced batch-to-batch variations, as the reaction conditions change with time, and may be difficult to scale-up, due to the changing power requirements for efficient fixing. They do, however, have a number of advantageous features. Primary amongst these is their simplicity both in use and in process development. For this reason they are preferred for small-scale production of highly priced products, especially where the same equipment is to be used for a number of different conversions. They offer a closely controllable environment that is useful for slow reactions, where the composition may be accurately monitored, and conditions (e.g. temperature, pH, coenzyme concentrations) varied throughout the reaction. They are also of use when continuous operation of a process proves to be difficult due to the viscous or intractable nature of the reaction mix.

Figure 5.1. Diagrams of various important enzyme reactor types.

a. Stirred tank batch reactor (STR), which contains all of the enzyme and substrates) until the conversion is complete;

b. batch membrane reactor (MR), where the enzyme is held within membrane tubes which allow the substrate to diffuse in and the product to diffuse out. This reactor may often be used in a semicontinuous manner, using the same enzyme solution for several batches;

c. packed bed reactor (PBR), also called plug -flow reactor (PFR), containing a settled bed of immobilised enzyme particles;

d. continuous flow stirred tank reactor (CSTR) which is a continuously operated version of (a);

e. continuous flow membrane reactor (CMR) which is a continuously operated version of (b);

f. fluidised bed reactor (FBR), where the flow of gas and/or substrate keeps the immobilised enzyme particles in a fluidised state.

All reactors would additionally have heating/cooling coils (interior in reactors (a), and (d), and exterior, generally, in reactors (b), (c), (e) and (f)) and the stirred reactors may

contain baffles in order to increase (reactors (a), (b), (d) and (e) or decrease (reactor (f)) the stirring efficiency. The continuous reactors ((c) -(f)) may all be used in a recycle mode where some, or most, of the product stream is mixed with the incoming substrate stream. All reactors may use immobilised enzymes. In addition, reactors (a), (b) and (e) (plus reactors (d) and (f), if semipermeable membranes are used on their outlets) may be used with the soluble enzyme.

The expected productivity of a batch reactor may be calculated by, assuming the validity of the non -reversible Michaelis -Menten reaction scheme with no diffusional control, inhibition or denaturation (see reaction scheme [1.7]   and equation (1.7). The rate of reaction (v) may be expressed in terms of the volume of substrate solution within the reactor (VolS) and the time (t):

            (5.1)

Therefore:

            (5.2)

On integrating using the boundary condition that [S] = [S]0 at time (t) = 0:

            (5.3)

Let the fractional conversion be X, where:

            (5.4)

Therefore;

            (5.4a)

and

            (5.4b)

Also

            (5.4c)

Therefore substituting using (5.4c) and (5.4b) in (5.3):

            (5.5)

The change in fractional conversion and concentrations of substrate and product with time in a batch reactor is shown in Figure 5.2(a).

Figure 5.2. This figure shows two related behaviours. 

(a) The change in substrate and product concentrations with time, in a batch reactor. The reaction S P is assumed, with the initial condition [S]0/Km = 10. The concentrations of substrate (??? and product (-----------) are both normalised with respect to [S]0. The normalised time (i.e. t? = t Vmax/[S]0) is relative to the time (t? = 1) that would

be required to convert all the substrate if the enzyme acted at Vmax throughout, the actual time for complete conversion being longer due to the reduction in the substrate concentration at the reaction progresses. The dashed line also indicates the variation of the fractional conversion (X) with t?. 

(b) The change in substrate and product concentrations with reactor length for a PBR. The reaction S P is assumed with the initial condition, [S]0/Km = 10. The concentrations of substrate (???) and product (-----------) are both normalised with respect to [S]". The normalised reactor length (i.e. I? = lVmax/F, where Vmax is the maximum velocity for unit reactor length and I is the reactor length) is relative to the length (i.e. when I? = 1) that contains sufficient enzyme to convert all the substrate at the given flow rate if the enzyme acted at its maximum velocity throughout; the actual reactor length necessary for complete conversion being longer due to the reduction in the substrate concentration as the reaction progresses. P may be considered as the relative position within a PBR or the reactor's absolute length.

Types of Reactors

        In an enzyme reactor, the highest specific enzyme activity is desirable.  It is considered an added bonus if the support that is used also aides in separation. One approach is to use a molecular sieve as the support and pulse the reactor bed with the alternating passage of substrate solution and water.  The result is that bands of unused substrate and product progress down the column. It so happens that the enzymes for which this technique would be useful are also those which in some cases benefit in having the enzyme immobilized on a porous support.

        For an industrial reactor, it is preferable to use supports that are non-biodegradable such as glass, silica, Celite, Bentonite, alumina, or titanium oxide, if possible. Even the linkages between enzyme and support can be non-biodegradable, as they are in the case of titanium. In some of these supports the physical nature of the surface becomes a major problem. Thus, some supports that form excellent packed beds fail to do so when coated with enzyme. Particles which ideally self-suspend in a fluid bed may form aggregates during use which will require more power to pump through substrate. Many problems were encountered using porous glass supports until someone realized that the glass itself could dissolve. This problem has been eliminated by treatment of the glass surface with zirconium.  

Many types of reactors have been proposed including the following: Batch reactors may include:

o Stirred Tank for Soluble Enzymes o Stirred Tank for Immobilized Enzymes o Stirred Tank with Immobilized Enzyme Basket Paddles o Stirred Tank with Immobilized Enzyme Basket Baffles o Total Recycle Packed Bed Reactor

o Total Recycle Fluidized Bed Reactor

 

Continuous reactors may include: o Stirred Tank Reactor with Filtration Recovery o Stirred Tank Reactor with Settling Tank Recovery o Stirred Tank Reactor with Immobilized Enzyme Basket Paddles o Stirred Tank Reactor with Ultra filtration Recovery o Packed Bed Reactor (same link as above) o Packed bed with recycle o Flat Bed Reactor o Filter Bed Reactor o Fluidized Bed Reactor, Same but better design (expanded top

section) o Membrane Reactor using hollow fibers

Combined CSTR/UF Reactor

       A combined CSTR/UF reactor is a combination of a continuously stirred tank reactor and an Ultra Filtration Unit. This type of reactor begins as a typical CSTR. However, the product passes through an Ultra Filtration Unit where the enzyme is removed and recycled back into the reactor. An example of what this combination rector can look like is shown below:

        In a combined CSTR/UF reactor the enzymes are immobilized in that they can not leave the reactor because of the filtration unit. This allows continuous processing with

free enzymes in the CSTR. The Ultra Filtration Unit contains a membrane which provides a semipermeable barrier which allows products and unreacted substrate, if there is any, to pass through while holding back the enzyme. There are other possibilities for similar reactors, such as combined reactor-separators. However, the combined CSTR/UF reactor has proven useful for several types of reactions where a typical immobilized enzyme would not be as effective. One example of this is for the conversion of benzylpenicillin to 6-aminopenicillanic acid by penicillin amidase.

  Packed Bed Reactor

        Continuous packed bed reactors are the most widely used reactors for immobilized enzymes and immobilized microbial cells. In these systems, it is necessary to consider the pressure drop across the packed bed or column, and the effect of the column dimensions on the reaction rate. There are three substrate flow possibilities in a packed bed and they are illustrated below:

1. Downward flow method 2. Upward flow method 3. Recycling method

        The recycling method is advantageous when the linear velocity of the substrate solution affects the reaction flow rate. This is because the recycling method allows the substrate solution to be passed through the column at a desired velocity.

        For industrial applications, upward flow is generally preferred over downward flow because it does not compress the beds in enzyme columns as downward flow does. When gas is produced during an enzyme reaction, upward flow is preferred.

        A continuous packed bed reactor has the following advantages over a batch packed bed reactor:

1. Easy, automatic control and operation 2. Reduction of labor costs 3. Stabilization of operating conditions 4. Easy quality control of products

   Recycle Reactor

        A recycle reactor is a reactor that is not seen very often, but is very important to consider when studying immobilized enzymes. It is very important to chemical engineering because it allows some substrates to be processed, which could not be processed using other reactor types. An example of a recycle reactor can be seen below:

        In a recycle reactor, a portion of the product stream is recycled and mixed with the inlet flow to the reactor. If the entire product stream is recycled back to the inlet stream, then it is called a total recycle reactor. This can obviously only be used in a batch process, because if the entire product stream is recycled back into the reactor in a continuous reactor, the volume of the reactor would increase to infinity. Therefore, we will only consider partial recycle streams in a continuous reactor on this page.

       This type of rector is used when you have a substrate that cannot be completely processed on a single pass, such as with an insoluble substrate. These reactors continue to move the same substrate through the reactor so that the effective contact time is high enough to allow the substrate to be processed. Recycle reactors also allow the reactor to operate at high fluid velocities. This is important because it minimizes the bulk mass

transfer resistance to the transport of the substrate. It is important to remember that a recycle reactor is simply a reactor, such as a CSTR or fluidized-bed reactor, with a recycle stream.

 Ultrafiltration Membrane Devices

        A continuous ultrafiltration membrane device is shown below:

        This device is suitable for a substrate of high molecular weight and a product of low molecular weight. Since the enzyme used here is soluble, no improvement in the stability of the enzyme can be expected. A hollow fiber device can also be used and its characteristics are essentially the same as those of an ultrafiltration membrane.