Crystals are grown in many shapes, which are dependent upon downstream processing or final product requirements.  {parse block="google_articles"}Crystal shapes can include cubic, tetragonal, orthorhombic, hexagonal, monoclinic, triclinic, and trigonal.  In order for crystallization to take place a solution must be "supersaturated".   Supersaturation refers to a state in which the liquid (solvent) contains more dissolved solids (solute) than can ordinarily be accomodated at that temperature.

Understanding the Basics of Crystallization

As with any separation method, equilibrium plays an important role.  Below is a general solubility curve for a solid that forms hydrate (a compound that has one or more water molecules attached) as it cools.

In Figure 1, X may be any solid that can form hydrates such as Na₂S₂O₃.  The number of hydrate molecules shown in Figure 1 is strictly arbitrary and will vary for each substance.

Figure 1: Example of a Solubility Curve

 So how do you grow crystals?  Let's consider an example that is fairly easy to envision.  Take a pot of boiling water and add table salt while stirring to make a water-salt solution.  Continue adding salt until no more salt will dissolve in the solution (this is a saturated solution).  Now add one final teaspoon of salt.  The salt that will not dissolve will help the first step in crystallization begin.  This first step is called "nucleation" or primary nucleation.  The salt resting at the bottom of the pot will provide a site for nucleation to occur.

On an industrial scale, a large supersaturation driving force is necessary to initiate primary nucleation.  The initiation of primary nucleation via this driving force is not fully understood which makes it difficult to model (experiments are the best guide).  Usually, the instantaneous formation of many nuclei can be observed "crashing out" of the solution.  You can think of the supersaturation driving force as being created by a combination of high solute concentration and rapid cooling.  In the salt example, cooling will be gradual so we need to provide a "seed" for the crystals to grow on.   In continuous crystallization, once primary nucleation has begun, the crystal size distribution begins to take shape.  Think about our salty water, as you look at Figure 2 describing the progression of crystallization.

 The second chief mechanism in crystallization is called secondary nucleation.  In this phase of crystallization, crystal growth is initiated with contact.  The contact can be between the solution and other crystals, a mixer blade, a pipe, a vessel wall, etc.  This phase of crystallization occurs at lower supersaturation (than primary nucleation) where crystal growth is optimal.   

Figure 2: Progression of Crystallization

Again, no complete theory is available to model secondary nucleation and it's behavior can only be anticipated by experimentation.  Mathematic relationships do exist to correlate experimental data.  However, correlating experimental data to model crystallization is time consuming and often considered extreme for batch operations, but can easily be justified for continuous processes where larger capital expenditures are necessary.  For batch operations, only preliminary data measurements are truly necessary.

We've discussed how crystallization occurs once supersaturation is reached, but how do we reach supersaturation?  We have already covered one such method in our salt crystallization example.  Since the solubility of salt in water decreases with decreasing temperature, as the solution cools, its saturation increases until it reaches supersaturation and crystallization begins (Figure 3).  Cooling is one of the four most common methods of achieving supersaturation.  It should be noted that cooling will only help reach supersaturation in systems where solubility and temperature are directly related.  Although this is nearly always the case, there are exceptions.  In Figure 3, you'll note that Ce₂(SO₄)₃ actually becomes less soluble in water at higher temperatures. 

 

Figure 3: Solubilities of Several Solids

The four most common methods of reaching supersaturation in industrial processes are:

1. Cooling (with some exceptions)
2. Solvent Evaporation
3. Drowning
4. Chemical Reaction

In an industrial setting, the solute-solvent mixture is commonly referred to as the "mother liquor".

Drowning describes the addition of a nonsolvent to the solution which decreases the solubility of the solid.  A chemical reaction can be used to alter the dissolved solid to decrease its solubility in the solvent, thus working toward supersaturation.  Each method of achieving supersaturation has its own benefits.  For cooling and evaporative crystallization, supersaturation can be generated near a heat transfer surface and usually at moderate rates.  Drowning or reactive crystallization allows for localized, rapid crystallization where the mixing mechanism can exert significant influence on the product characteristics.

Equipment Used in Crystallization

Tank Crystallizers

This is probably the oldest and most basic method of crystallization.  In fact, the "pot of salt water" is a good example of tank crystallization.  Hot, saturated solutions are allowed to cool in open tanks.  {parse block="google_articles"}After crystallization, the mother liquor is drained and the crystals are collected.  Controlling nucleation and the size of the crystals is difficult.   The crystallization is essentially just "allowed to happen".  Heat transfer coils and agitation can be used.  Labor costs are high, thus this type of crystallization is typically used only in the fine chemical or pharmaceutical industries where the product value and preservation can justify the high operating costs.

Scraped-Surface Crystallizers

An example may be the Swenson-Walker crystallizer consisting of a trough about 2 feet wide with a semi-circular bottom.  The outside is jacketed with cooling coils and an agitator blade gently passes close to the trough wall removing crystals that grow on the vessel wall.

Forced Circulating Liquid Evaporator-Crystallizer

Just as the name implies, these crystallizers combine crystallization and evaporation, thus the driving forces toward supersaturation.  The circulating liquid is forced through the tubeside of a steam heater.  The heated liquid flows into the vapor space of the crystallization vessel.  Here, flash evaporation occurs, reducing the amount of solvent in the solution (increasing solute concentration), thus driving the mother liquor towards supersaturation.  The supersaturated liquor flows down through a tube, then up through a fluidized area of crystals and liquor where crystallization takes place via secondary nucleation.  Larger product crystals are withdrawn while the liquor is recycled, mixed with the feed, and reheated.

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Figure  4: Forced Circulating Liquid Evaporator-Crystallizer

Circulating Magma Vacuum Crystallizer

In this type of crystallizer, the crystal/solution mixture (magma) is circulated out of the vessel body.  The magma is heated gently and mixed back into the vessel.  A vacuum in the vapor space causes boiling at the surface of the liquid.  The evaporation causes crystallization and the crystals are drawn off near the bottom of the vessel body.

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Figure 5: Circulating Magma Vacuum Crystallizer

References

1. Price, Chris J., "Take Some Solid Steps to Improve Crystallization", Chemical Engineering Progress, September 1997, p. 34.
2. Geankoplis, Christie J., Transport Processes and Unit Operations, 3rd Ed., Prentice Hall, New Jersey, 1993, ISBN: 0-13-930439-8
3. Brown, Theodore L., Chemistry: The Central Science, 5th Ed., Prentice Hall, New Jersey, 1991, ISBN: 0-13-126202-5
4. Swenson Process Equipment web site at http://www.swensontechnology.com/

 

 

Both equipment and utilities costs increase markedly and the operation can become uneconomical. If the system forms azeotropes, as in a benzene and cyclohexane system, a different problem arises, - the azeotropic composition limit the separation, and for a better separation this azeotrope must be bypassed in some way. At low to moderate pressure, with the assumption of ideal-gas model for the vapor phase, the vapor-liquid phase equilibrium (VLE) of many mixture can be adequately describe by the following Modified Raoult's Law:

Eq. (1)

 where:{parse block="google_articles"}

y_i = mole fraction of component i in vapor phase
x_i = mole fraction of component i in liquid phase
P = system pressure
P^sat = vapor pressure of component i
?_i = liquid phase activity coefficient of component i

When ?_i = 1, the mixture is said to be ideal Equation 1 simplifies to Raoult's Law. Nonideal mixtures   (?_i ? 1) can exhibit either positive (?_i > 1) or negative deviations (γ_i < 1) from Raoult's Law. In many highly nonideal mixtures these deviations become so large that the pressure-composition (P-x, y) and temperature-composition (T-x, y) phase diagrams exhibit a minimum or maximum azeotrope point. In the context of the T-x, y phase diagram, these points are called the minimum boiling azeotrope (where the boiling temperature of the azeotrope is less than that of the pure component) or maximum boiling azeotrope (the boiling temperature of the azeotrope is higher than that of the pure components). About 90% of the known azeotropes are of the minimum variety. At these minimum and maximum boiling azeotrope, the liquid phase and its equilibrium vapor phase have the same composition, i.e.,

Two main types of azeotropes exist, i.e. the homogeneous azeotrope, where a single liquid phase is in the equilibrium with a vapor phase; and the heterogeneous azeotropes, where the overall liquid composition which form two liquid phases, is identical to the vapor composition.

Most methods of distilling azeotropes and low relative volatility mixtures rely on the addition of specially chosen chemicals to facilitate the separation. The selection of the separating agent will be discussed later.

The five methods for separating azeotropic mixtures are:

Residue Curve Maps

The most simple form of distillation, called simple distillation, is a process in which a muticomponent liquid mixture is slowly boiled in an open pot and the vapors are continuously removed as they form. At any instant in time the vapor is in equilibrium with the {parse block="google_articles"}liquid remaining on the still. Because the vapor is always richer in the more volatile components than the liquid, the liquid composition changes continuously with time, becoming more and more concentrated in the least volatile species. A simple distillation residue curve is a graph showing how the composition of the liquid residue curves on the pot changes over time. A residue curve map is a collection of the liquid residue curves originating from different initial compositions. Residue curve maps contain the same information as phase diagrams, but represent this information in a way that is more useful for understanding how to synthesize a distillation sequence to separate a mixture.

All of the residue curves originate at the light (lowest boiling) pure component in a region, move towards the intermediate boiling component, and end at the heavy (highest boiling) pure component in the same region. The lowest temperature nodes are termed as unstable nodes (UN), as all trajectories leave from them; while the highest temperature points in the region are termed stable nodes (SN), as all trajectories ultimately reach them. The point that the trajectories approach from one direction and end in a different direction (as always is the point of intermediate boiling component) are termed saddle point (S). Residue curve that divide the composition space into different distillation regions are called distillation boundaries. A better understanding of the residue curve map may be view in Figure 1.  Notice that the trajectories move from the lowest temperature component towards the highest.

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Figure 1: Residue curve map for a ternary mixture with a distillation boundary running from pure component D to the binary azeotrope C.

Residue curve maps would be of limited usefulness if they could only be generated experimentally. Fortunately that is not the case. Using various references, the simple distillation process can be described by the set of equations:

Eq. (3)

 where:

y_i = mole fraction of component i in vapor phase
x_i = mole fraction of component i in liquid phase
ξ = nonlinear time scale
c = number of component in the mixture

Research studies have also been done to determine the relationship between the number of nodes (stable and unstable) and saddle points one can have in a legitimately drawn ternary residue plot. The equation is based on topological arguments. One form for this equation is:

where:
N_i = number of nodes (stable and unstable) involving i species
S_i = number of saddles involving i species

Many different residue curve maps are possible when azeotropes are present. Ternary mixtures containing only one azeotrope may exhibit six possible residue curve maps that differ by the binary pair forming the azeotrope and by whether the azeotrope is minimum or maximum boiling.

Even though the simple distillation process has no practical use as a method for separating mixtures, simple distillation residue curve maps have extremely useful applications, such as:

1. Testing of the consistency of experimental azeotropic data;
2. Predict the order and content of the cuts in batch distillation;
3. In distillation, to check whether the given mixture is separable by distillation, identification of the entrainers / solvents, prediction of attainable product compositions, qualitative prediction of composition profile shape, and synthesis of the corresponding distillation column.

By identifying the limiting separation achievable by distillation, residue curve maps are also useful in synthesizing separation sequences combining distillation with other methods.

Homogenous Azeotropic Distillation

The most general definition of homogeneous azeotropic distillation is the separation of any single liquid-phase mixture containing one or more azeotropes into the desired pure component or azeotropic products by continuous distillation. Thus, in addition to azeotropic mixtures which require the addition of a miscible separating agent in order to be separated, homogeneous azeotropic distillation also includes self-entrained mixtures that can be separated without the addition of a separating agent.{parse block="google_articles"}

The first step in the synthesis of a homogeneous azeotropic distillation sequence is to determine the separation objective. Sometimes it is desirable to recover all of the constituents in the mixture as pure components other times it is sufficient to recover only some of the pure components as product. In our example, we would like to recover the cyclohexane product at 90% purity and recycle the separating agent back to the initial separating column for further use.

The second step is to sketch the residue curve map for the mixture to be separated. The residue curve map allows one to determine whether the goal can be reached and if so how to reach it, or the goal needs to be redefined.

Distillation boundaries for continuous distillation are approximated by simple distillation boundaries. It is a good approximation for mixtures with nearly simple distillation boundaries. For a separation to be feasible by distillation, the simple distillation boundary should not be crossed, i.e. the distillate and bottom composition should lie in the same distillation region. A more detail calculation method involving the composition will be discuss in the later section.

In the most common situation, a separating agent is added to separate a minimum boiling binary azeotrope into its two constituent pure components by homogeneous azeotropic distillation. Michael F. D. and Jeffrey P. K. presented seven of the most favorable residue curve maps for this task. Of the seven, the map representing extractive distillation is by far the most common and the most important. Its corresponding residue curve and column sequences are shown in Figure 2 below.

Extractive Distillation

Extractive distillation is defined as distillation in the present of a miscible, high boiling, relatively nonvolatile component, the solvent, that forms no azeotropes with the other components in the mixture. It is widely used in the chemical and petrochemical industries for separating azeotropic, close-boiling, and others low relative volatility mixture.

Figure 2: Extractive Distillation with a Heavy Solvent

Extractive distillation works because the solvent is specially chosen to interact differently with the components of the original mixture, thereby altering their relative volatilities. Because these interactions occur predominantly in the liquid phase , the solvent is continuously added near the top of the extractive distillation column so that an appreciable amount is present in the liquid phase on all of the trays below. The mixture to be separated is added through second feed point further down the column. In the extractive column, the component having the greater volatility, not necessarily the component having the lowest boiling point, is taken overhead as a relatively pure distillate. The other component leaves with the solvent via the column bottoms. The solvent is separated from the remaining components in a second distillation column and then recycled back to the first column.

There are several industrial application for homogeneous azeotropic distillation listed in the Encyclopedia of Separation Technology by Michael F. D., Jeffrey P. K.

Extractive distillations can be divided into three categories, each correspond to the different residue curve maps, the minimum boiling azeotropes, maximum boiling azeotropes and the nonazeotrope mixtures. Since our benzene-cyclohexane mixture to be separated is of the second type of mixture, i.e. the minimum boiling azeotrope, we will focus our attention on column sequencing this type of azeotropic separation method in the following section.

As in azeotropic distillation, design of extractive distillation system will also requires significant preliminary work including:

• Choosing the solvent
• Developing or finding necessary data, such as azeotropic condition or residue curve
• Preliminary screening
• Computer simulation
• Small scale testing

For our example, we will consider the first four steps.

Solvent Screening and Selection

Solvent Criteria

One of the most important steps in developing a successful (economical) extractive distillation sequence is selecting a good solvent. Approaches to the selection of an extractive distillation solvent are discussed by Berg, Ewell et al. ^, and Tassions. In general, selection criteria include the following :

1. Should enhance significantly the natural relative volatility of the key component. {parse block="google_articles"}
2. Should not require an excessive ratio of solvent to nonsolvent (because of cost of handling in the column and auxiliary equipment.
3. Should remain soluble in the feed components and should not lead to the formation of two phase.
4. Should be easily separable from the bottom product.
5. Should be inexpensive and readily available.
6. Should be stable at the temperature of the distillation and solvent separation.
7. Should be nonreactive with the components in the feed mixture.
8. Should have a low latent heat.
9. Should be noncorrosive and nontoxic.

Naturally no single solvent or solvent mixture satisfy all the criteria, and compromises must be reached.

Solvent Screening

Perry's handbook serve as a good reference for the solvent selection procedure, which can be thought of as a two-step process, i.e.:

Broad screening by functional group or chemical family

1. Homologous series : Select candidate solvent from the high boiling homologous series of both light and heavy key components.  
2. Robins Chart: Select candidate solvents from groups in the Robbins Chart (part of the chart is shown in Table 3) that tend to give positive (or no) deviations from Raoult's law for the key component desire in the distillate and negative (or no) deviations for the other key.
3. Hydrogen-bonding characteristic: are likely to cause the formation of hydrogen bonds with the key component to be removed in the bottoms, or disruption of hydrogen bonds with the key to be removed in the distillate. Formation and disruption of hydrogen bonds are often associated with strong negative and positive deviations, respectively from Raoult's Law.
4. Polarity characteristic: Select candidate solvents from chemical groups that tend to show higher polarity than one key component or lower polarity than the other key.

Identification of individual candidate solvents

1. Boiling point characteristic: Select only candidate solvents that boil at least 30-40^oC above the key components to ensure that the solvent is relatively nonvolatile and remains largely in the liquid phase. With this boiling point difference, the solvent should also not form azeotropes with the other components.
2. Selectivity at the infinite dilution: Rank the candidate solvents according to their selectivity at infinite dilution.
3. Experimental measurement of relative volatility: Rank the candidate solvents by the increase in relative volatility caused by the addition of the solvent.

Residue curve maps are of limited usefulness at the preliminary screening stage because there is usually insufficient information available to sketch the them, but they are valuable and should be sketched or calculated as part of the second stage of the solvent selection.

A Sample Distillation Process

For our example which deals with the azeotropic mixture formed between benzene and cyclohexane, we have chosen extractive distillation (one of the homogeneous azeotropic distillation methods). The reason of choosing this method is due to the availability of information regarding this separation technique and its tendency to operate more efficiently, i.e. in separating and recycling the separating agent. A brief discussion of the process is given below.{parse block="google_articles"}

After the mixture exited as the bottom product of the flash unit, it contains mostly our desire product of cyclohexane and also a significant amount of unreacted benzene, which is to be recycled back to the reactor for further conversion. Our main goal now is to further separate the remaining components in the mixture. As cyclohexane and benzene have been encounter most of the remaining composition with the mole % of 44.86 and 54.848 respectively (Table 1), we will consider this to be a binary mixture in our further discussion.

From the process flowsheeting, we would like to operate the distillation column at the pressure of 150 kPa. At this condition, cyclohexane and benzene will have boiling points of 94.34 ^°C and 93.49 ^°C respectively (Figure 3). This is a typical case where conventional distillation would struggle to perform the separation of this type of close boiling mixture. Thus, a special type of distillation technique, i.e. extractive distillation has been chosen in order to purify the desire product, i.e. cyclohexane to our desired purity of 99.3%.

As can be shown from Figure 3, this binary composition will form a minimum boiling, homogeneous azeotrope at the temperature of 91^oC and the corresponding composition at this point will be 45.5 mole % for cyclohexane and 55.5 mole % for benzene (Figure 4).

 

Figure 3: T-xy Plot for Benzene and Cyclohexane	Figure 4: x-y Plot for Benzene and Cyclohexane at 150 kPa

 

Solvent Selection for the Benzene-Cyclohexane Binary Mixture

In order to perform a successive extractive distillation, a solvent needs to be chosen to "break" the azeotrope that forms at the operating pressure of the distillation column. Recommended solvent for the benzene-cyclohaxane mixture from the literature^,,, is aniline, with a solvent to feed ratio (S/F) of 4, which will shift the azeotropic point toward the corner of the high-boiling component cyclohexane, and the equilibrium curve of the original components fall below the diagonal (Figure 5).

Figure 5: Elimination of Azeotropic Point with the Addition of a Separating Agent

As was stated in the above section, the primary goal of solvent selection is to identify a group of feasible solvents to perform a good separation. The desired product, i.e. cyclohexane should have a purity of above 99% to meet the market standard. Aniline was the first solvent that had been put to the simulator to be tried out, as it is of the same homologous group as benzene. As can be shown from the result in Table 2, this solvent will produce the desire production rate of 150 with the solvent flow rate of 3500, i.e. a S/F ratio of 9.85. However, the product purity can only reach 70.08% and this does not meet our product specification. As a result, other solvent may have to be researched to perform the desire separation.

We will have to perform the solvent selection criteria as stated in the preceding section. At the column pressure of 150 kPa, cyclohexane and benzene boil at 94.34 ^°C and 93.49 ^°C respectively and form a minimum-boiling azeotrope at 91 ^°C. The natural volatility of the system is benzene > cyclohexane, so the favored solvents most likely will be those that cause the benzene to be recovered in the distillate. However, in order to get a better quality of product, we would like to recover cyclohexane as the distillate rather than from the bottom stream. Thus, solvent to be chosen should give positive deviations from Raoult's law for cyclohexane and negative (or zero) deviation for benzene.

 

Turning to the Robbins Chart (Table 3), we note that solvents that may cause the positive deviation for cyclohexane (Class 12) and negative (or zero) to benzene (Class 11) came from the groups of 4, 7, 8 and 9, which consist of polyol, amine and ether. We further consider the solubility, the hydrogen bonding effect, and also the homologous characteristic of the solvent with the corresponding components in the feed mixture. As few candidate solvents that had been put to the computer simulation, included phenol (homologous to benzene), 1,2-benzenediol (homologous to benzene, with -OH group that will produce hydrogen bonding), 1,3-butanediol (with -OH group that will produce hydrogen bonding), and also 1,2-propanediol (same characteristic as with 1,3-butanediol). 1,2-propanediol (often known as propylene glycol), seem to give the most promising results compared to the other solvents. This result may be caused from the high solubility of benzene in this solvent and the hydrogen bonding that were formed between the two constituents. Simulation result of this solvent can be view in Table 4.

Construction of the Residue Curve

Equation 3 and 4 were used to sketch the corresponding residue curve for the three species. From the above information, we know that these species have boiling points at 94.34 (cyclohexane), 93.49 (benzene) and 200.35 ^oC (propylene glycol) at the pressure of 150 kPa, and an azeotrope that boils at 91 ^°C between the two more volatile species. As were shown from Figure 6 and Figure 7 there were no new azeotropes formed between the solvent 1,2-propanediol respectively with the another two component in the feed.

 

We then start to sketch our residue curve map by sketching the triangular diagram in Figure 8, and placing the arrows pointing from the lower to higher temperatures around the edge. The corner points for benzene and cyclohexane are single species point, and both are unstable nodes - all residue curves leave. The corner point for propylene glycol is a single species point which is a stable node - all residue curve enter. All three are nodes; none are saddles, thus;

N₁ = 3 and S₁ = 0

We then further assume that there will be no ternary azeotrope been form among the three constituents, i.e.,

N₃ = S₃ = 0 

Figure 6: x-y Plot for Benzene and Propylene Glycol	Figure 7: x-y Plot for Cyclohexane and Propylene Glycol

 

Figure 8: Preliminary Sketch for Residue Curve

The remaining steps here require the identification of the only binary azeotrope that form between benzene and cyclohexane, to be either a node or a saddle point. From equation 4:

4(0-0) + 2(N₂ - S₂) + (3-0) = 1
2(N₂ - S₂) = -2
N₂ - S₂ = -1

Thus, the only way we can satisfy the above equation is letting N₂ = 0 and S₂ = 1, i.e. the binary azeotrope is a saddle point, which directs the trajectories in another direction.

 

Figure 9: Completed Sketch for Residue Curve

Column Operation

The extractive distillation unit of this cyclohexane production plant consists of two distillation columns (Figure 10), which we can easily classify as direct sequence columns. The first column acts as an extractive column where the solvent is introduced at the second stage of the column, so that it will be present throughout the column and exits with the bottoms. As were stated above, the solvent alters the natural volatility of the binary mixture by forming hydrogen bonds with benzene and allowing it to be recovered as the bottom product.

The bottom product of the first column will then fed to the second column, i.e. the solvent recovery column, to undergo the normal distillation to separate both the components for further usage, i.e. benzene being recycled to the reactor for further conversion while solvent to the first column for reuse. The main operation parameter of the distillation unit is shown in Table 5.

 

Figure 10: Extractive Distillation Unit for Cyclohexane Production Plant

 

 

Figure 11: Distillation Ternary Diagram for the Extractive Distillation Unit

 

References

1. J. M. Smith, H. C. Van Ness, M. M. Abbott, 1996, Introduction to Chemical Engineering Thermodynamics, McGraw Hill, p.449.
2. Douglas M. Ruthven (Ed.), 1997, Encyclopedia of Separation Technology, Vol. 1, John Wiley, [Distillation, Azeotropic and Extractive by Michael F. D., Jeffrey P. K.]{parse block="google_articles"}
3. L. T. Biegler, E. I. Grossmann, Arthur W. Westerberg, 1997, Systematic Methods of Chemical Process Design, Prentice Hall, [Separating Azeotropic Mixture]
4. Green, Perry, 1997, Perry's Chemical Engineers' Handbook, 7^th Edition, McGraw Hill,. [Section 13-78 Enhanced Distillation by J. D. Seader, Jeffrey J. S., Scott D. B.]
5. Philip A. Schweitzer, 1988, Handbook of Separation Technique for Chemical Engineers, 2^nd Edition, McGraw Hill, [Continuous Distillation: Separation of Multicomponent Mixture by Edward C. R., John E. M.]
6. James R. F., Distillation.
7. Robert F. G., 1972, Extractive and Azeotropic Distillation, [Rapid screening of Extractive Distillation Solvent. Predictive and Experimental Techniques by Tassios P. D.]
8. Green, Perry, 1997, Perry's Chemical Engineers' Handbook, 7^th Edition, McGraw Hill,. [Section 13-79 Enhanced Distillation by J. D. Seader, Jeffrey J. S., Scott D. B.]
9. James R. Fair, Distillation [Special Distillation]
10. Kith-Othmer, 1965, Encyclopedia of Chemical Technology, 2^nd Edition, Vol. 6, John Wiley, [Cyclohexane by James J.Kirk]
11. Lee L. E., 1999, Plant Design Project Cyclohexane Production, UTM, {Chapter 7 Discussion].

The following need to be carefully evaluated when optimizing the design and operation of the extraction processes.{parse block="google_articles"}

• Solvent selection
• Operating Conditions
• Mode of Operation
• Extractor Type
• Design Criteria

Solvent Selection

Solvents differ in their extraction capabilities depending on their own and the solute's chemical structure. Robbins (1) presents a table showing Organic-Group interactions from which one can identify the desired functional group(s) in the solvent for any given solute.

Once the functional group is identified, possible solvents can be screened in the laboratory. The distribution coefficient and selectivity are the most important parameters that govern solvent selection. The distribution coefficient (m) or partition coefficient for a component (A) is defined as the ratio of concentration of a A in extract phase to that in raffinate phase. Selectivity can be defined as the ability of the solvent to pick up the desired component in the feed as compared to other components. The desired properties of solvents are a high distribution coefficient, good selectivity towards solute and little or no miscibility with feed solution. Also, the solvent should be easily recoverable for recycle. Designing an extractor is usually a fine balance between capital and operating costs. Usually, good solvents also exhibit some miscibility with feed solution (see Table 1). Consequently, while extracting larger quantities of solute, the solvent could also extract significant amount of feed solution.

Other factors affecting solvent selection are boiling point, density, interfacial tension, viscosity, corrosiveness, flammability, toxicity, stability, compatibility with product, availability and cost.

For an existing process, replacing the solvent is usually a last resort because this this would call for going back to laboratory screening of the solvent and process optimization. However, changes in environmental regulations and economic considerations often induce the need to improve the processes in terms of solute recovery. Also the availability of specialized and proprietary solvents that score over conventional solvents in terms of performance and economics for several extraction processes can provide additional incentives for a solvent change.

Selection of Extraction Conditions

Depending on the nature of the extraction process, the temperature, pH and residence time could have an effect on the yield and selectivity. Operating pressure has a negligible affect on extraction performance and therefore most extractions take place at atmospheric pressure unless governed by vapor pressure considerations.

Temperature can also be used as a variable to alter selectivity. Elevated temperatures are sometimes used in order to keep viscosity low and thereby minimizing mass-transfer resistance. Other parameters to be considered are selectivity, mutual solubility, precipitation of solids and vapor pressure.

The pH becomes significant in metal and bio-extractions. In bio-extractions (e.g., Penicillin) and some agrochemicals (e.g. Orthene), pH is maintained to improve distribution coefficient and minimize degradation of product. In metal extractions, kinetic considerations govern the pH. In dissociation-based extraction of organic molecules, pH can play a significant role (e.g., cresols separation). Sometimes, the solvent itself may participate in undesirable reactions under certain pH conditions (e.g., ethyl acetate may undergo hydrolysis in presence of mineral acids to acetic acid and ethanol).

Residence time is an important parameter in reactive extraction processes (e.g., metals separations, formaldehyde extraction from aqueous streams) and in processes involving short-life components (e.g., antibiotics & vitamins)

Selection of Mode of Operation

Extractors can be operated in crosscurrent or counter-current mode. The following section compares these configurations.

Cross-Current Operation

Crosscurrent mode is mostly used in batch operation. Batch extractors have traditionally been used in low capacity multi-product plants such as are typical in the pharmaceutical and agrochemical industries. For washing and neutralization operations that require very few stages, crosscurrent operation is particularly practical and economical and offers a great deal of flexibility. The extraction equipment is usually an agitated tank that may also be used for the reaction steps. In these tanks, solvent is first added to the feed, the contents are mixed, settled and then separated. Single stage extraction is used when the extraction is fairly simple and can be achieved without a high amount of solvent. If more than one stage is required, multiple solvent-washes are given.{parse block="google_articles"}

Though operation in crosscurrent mode offers more flexibility, it is not very desirable due to the high solvent requirements and low extraction yields. The following illustration gives a quick method to calculate solvent requirements for crosscurrent mode of extraction.

A single-stage extractor can be represented as:

where

F = Feed quantity / rate, massR = Raffinate quantity / rate, massS = Solvent quantity / rate, massE = Extract quantity / rate, mass

X_f, X_r, Y_s, and Y_e are the weight fractions of solute in the feed, raffinate, solvent and extract, respectively.

Partition coefficient 'm' is defined as the ratio of Y_e to X_r at equilibrium conditions

The flows and concentrations are represented in solute-free basis as such a representation leads to simplification of equations. For example, for a 100 kg/hr feed containing 10% weight acetic acid, F = 100-10 = 90 kg/hr, X_r = 0.1/(1-0.1) = 0.111

The component mass balance can be represented as:F X_f + S Y_s = R X_r + E Y_e

Assuming (i) immiscibility of feed and solvent and (ii) the initial solvent is free of solute, i.e., F = R, S = E and Y_s = 0 and using the equilibrium relation of Y_e = m X_r, this equation simplifies to

S = F/m (X_f /X_r-1)

or

reduction ratio, X_f /X_r = 1+ m S/F

For multi-stage crosscurrent operation:

Assuming that the partition coefficient (m) is constant over the concentration range and the solvent quantity in each of the 'n' stages is the same, i.e., S₁ = S₂ =.....=S _n = S/n,

Solvent Requirement is

S = n * F/m [(X_f /X_r)^1/n - 1]

reduction ratio X_f /X_r = (1+mS/nF)ⁿ

It can be proved mathematically that the total solvent quantity would be minimum if the solvent were distributed equally between washes.

Diminishing Returns

The following chart shows solvent requirements for a typical reduction ratio (X _f /X_r) of 10 using crosscurrent extraction.

Figure 1: Reduction Ratio for Crosscurrent Extraction

With one stage, 18,000 kg of solvent is required for 1,000 kg of feed (m = 1 and X_f / X_r = 10). With two stages, solvent requirement reduces to 8,650 kg, and with three stages, it reduces further to 6,930 kg. However, as can be seen from the chart, using more than three stages has minimal effect on solvent usage. This fact combined with practical limitations of solvent handling and increased batch time confines the number of solvent washes to three.

Counter-Current Operation

As described above, the crosscurrent operation is mostly used in low capacity multi-product batch plants. For larger volume operation and more efficient use of solvent, countercurrent mixer-settlers or columns are employed. Countercurrent operation conserves the mass transfer driving force and hence gives optimal performance.

Equations for countercurrent extraction get more complicated with increasing number of stages. It can be shown that for a 'n' stage operation, the raffinate concentration would be

X_r = X_f * (mS/F - 1)/ ( [mS/F] ⁿ⁺¹ -1)

The solvent requirement for any raffinate concentration X_r could be determined by iteration from the above equation.

For mS/F = 1, the equation takes the form of X_r = X_f / (n + 1)

The dimensionless term mS/F, included in all the above equations, is called the extraction factor (E), and is an important parameter in the design of extraction processes. For a given number of stages, the higher the E factor, the higher is the reduction ratio and easier is the extraction. Systems with E of less than 1.3 are not likely to be commercially feasible.

Comparing Counter versus Cross Flow Operation

The following graph compares the reduction ratios (X_f / X_r) of the crosscurrent and countercurrent modes of operation.

Figure 2: Comparing Counter and Cross Flow Extraction

The graph shows that for a given extraction factor (E), and number of stager (n), the countercurrent mode of operation outperforms the crosscurrent mode. This is demonstrated in a case study presented at the end of this paper. Koch-Glitsch has demonstrated these benefits on their pilot and commercial scale extraction columns for several systems.

The equations given above can be used to compare solvent requirements for various modes of operation and can serve as a starting point for identifying scope for optimizing solvent quantity. However, these equations should be used with caution as the assumptions of immiscibility, constancy of partition coefficient over desired range and solute-free fresh solvent are not valid in all practical applications.

As the solvent quantity is reduced, the solute concentration in the extract increases. This usually affects the physical properties and the selectivity. Therefore optimization exercise should be backed up by laboratory extraction data.

Selecting the Type of Extractor

Commercially important extractors can be classified into the following broad categories.

Mixer-Settlers

Centrifugal devices

Column contactors (static) - Examples include spray columns, sieve plate columns, and random or structurally packed columns.

Column contactors (agitated) - Agitated columns can be further split into rotary or reciprocating type. Examples of rotary agitated columns include rotary disc contactor, Scheibel column, Kuhni column. Examples of reciprocating agitated columns include the Karr column and the pulsed column.

Mixer-Settlers

As the name indicates, are usually a series of static or agitated mixers interspersed with settling stages. These are mostly used in the metal industry where intense mixing and high residence time is required by the reactive extraction processes.

In batch mode of operation, these mixer-settlers could be simple batch vessels where feed and solvent are mixed and settled. This operation is repeated with fresh solvent washes as described earlier.

Centrifugal Contactors

Centrifugal contactors are high-speed rotary machines that offer advantages of very low residence time. The number of stages in a centrifugal device is usually limited to one, but currently devices with multiple numbers of stages are common. These extractors are mainly used in pharmaceutical industry.

Counter-current Column Contactors

Counter-current column contactors are most popular in the chemical industry. These could be static or agitated. Several types of extractors are available (see Table 2) and each has its own advantages.

Factors Affecting the Selection of Contactors

Important factors to consider when selecting extractor types are the stage requirements, fluid properties and operational considerations. The following table outlines the capabilities and characteristics of different extractor-types:

The Karr reciprocating plate extractor can effectively handle low interfacial tension systems. Other factors governing extractor selection are presence of solids, safety and maintenance requirements.

Design Criteria

The basic function of extraction equipment is to mix two phases, form and maintain droplets of dispersed phase and subsequently separate the phases. The following section outlines some of the factors that need to be considered while designing and optimizing extraction equipment.

1. Mixing - The amount of mixing required is determined by physical properties such as viscosity, interfacial tension and density differences between the two phases. It is important to provide just the right amount of mixing. Less mixing causes the formation of large droplets and decreases interfacial area (interfacial area varies with the square of the droplet diameter). This reduces mass transfer and decreases stage efficiency. Higher agitation (more mixing) minimizes mass transfer resistance during reactions and extraction but contributes to the formation of small and difficult-to-settle droplets or emulsions. In agitated batch extractors, the agitator design is often optimized for reaction and heat transfer, not extraction, as these are generally multi-purpose vessels.The agitator imparts maximum energy at the tip where the velocity is highest and minimum energy at the center. This creates non-uniform droplet sizes, with the smallest being formed at the agitator tip. Reaching extraction equilibrium is controlled by the largest droplet size and the smallest droplet controls settling time. Therefore, over-agitation sometimes takes its toll by causing difficulties in phase separation. Usually a redesign in terms of configuration or change in agitation speed helps in optimizing batch time.Static extraction columns rely completely on the packing/internals and fluid flow velocities past the internals to create turbulence and droplets. Therefore these are restricted by minimum flow requirement of at least one of the phases. Agitated columns have more operating flexibility as the specific energy input can be varied in most designs. Axial mixing (along column length) in column contactors reduces stage efficiency. Baffles or similar arrangements are used to minimize axial mixing in static as well as agitated columns. It is also important to avoid temperature gradients in columns to prevent thermal currents contributing to axial mixing.
2. Settling - The settling characteristics depend on the fluid properties (density difference, interfacial tension, and continuous phase viscosity) and the amount of mixing. Settling in agitated batch vessels is carried out by stopping the agitator. In continuous columns, a settling section is provided either as a part of the extractor or as a separate piece of equipment downstream of the extractor. Emulsions are usually formed due to over agitation and in such cases, settling needs to be carried out over an extended period. Emulsions can also form due to the inherent nature of the chemical compounds involved or due to contaminants that substantially lower the interfacial tension. Sometimes coagulants are added to prevent or minimize emulsification. Passing the emulsion layer through a coalescer can break some of these emulsions. In continuous extractors, the creation of emulsions is less severe as good droplet size distribution can be attained at lower agitation speeds in a lesser diameter column. Also, columns such as the Karr reciprocating plate extractor impart uniform energy throughout the radius as a result of the reciprocating motion and this creates a much narrower droplet distribution.A similar phenomenon to emulsions is the formation of a 'rag layer'. This is a layer containing loose solid substances that float at the interface. These solid substances are generally foreign impurities that exist in the feed streams or those that precipitate from the system during extraction. In continuous extraction the liquid interface containing the rag layer can be continuously withdrawn, filtered and sent back to extractor. Selection of continuous and dispersed phases can have an effect on formation of emulsion and rag layer. Reversing continuous and dispersed phases sometimes drastically reduces or eliminates emulsion formation. Changing extraction temperature could also help in reducing emulsion and rag layer.
3. Selection of Continuous and Dispersed Phase - In column extractors, the phase with the lower viscosity (lower flow resistance) is generally chosen as the continuous phase. Also note that the phase with the higher flow rate can be dispersed to create more interfacial area and turbulence. This is accomplished by selecting an appropriate material of construction with the desired wetting characteristics. In general, aqueous phases wet metal surfaces and organic phases wet non-metallic surfaces. Change in flows and physical properties along the length of extractor should also be considered. Choosing a continuous phase is generally not available in batch processes, as the larger liquid phase usually becomes the continuous phase.

Conclusions

As we have seen in the previous sections, there are a number of factors affecting extraction performance. Laboratory and pilot plant testing using actual feed and solvent help immeasurably in optimization. The study could often be an iterative cycle involving laboratory testing followed by process simulation and design. In most industrial extractors, there is usually a good scope for optimizing solvent usage and energy consumption.

References

1. Robbins, Chem. Eng. Prog., 76(10), 58-61 (1980).
2. Cusack, R.W., & Glatz, D., et al, "A Fresh Look at Liquid-Liquid Extraction", Chemical Engineering, February, March & April 1991.

{parse block="google_articles"}The excellent mass-transfer properties conferred by the hollow fiber configuration soon led to numerous commercial applications in various field such as the medical field (blood fractionation), water reclamation (purification and desalination), gas separation, azeotropic mixture separation (using pervaporation). Others application of this type of membrane are in various stage of development, e.g. and the biochemical industry (bioseparation and bioreactor) and hydrocarbon separation (by pervaporation). Due to the high technology of this advanced materials, Malaysia is urged to focus more in the research of this materials as it has showed its distinguish performance in various field and applications compared to the conventional technique.

Historical Development of Membranes

Table 1 below indicates some of the historical development of the membrane technology before the "Golden Age" of membrane technology:

The "Golden Age" of membrane technology (1960-1980) began in 1960 with the invention by Loeb and Sourirajan of the first asymmetric integrally skinned cellulose acetate RO membrane^,. This development simulated both commercial and academic interest, first in desalination by reverse osmosis, and then in other membrane application and processes. During this period, significant progress was made in virtually every phase of membrane technology: applications, research tools, membrane formation processes, chemical and physical structures, configurations and packaging.

Basic Morphology

Figure 1: Membrane Morphology

Two basic morphology of hollow fiber membrane are isotropic and anisotropic (Figure 1). Membrane separation is achieved by using of this morphologies. The anisotropic configuration is of special value. In the early 1960s, the development of anisotropic membranes exhibiting a dense, ultrathin skin on a porous structure provided a momentum to the progress of membrane separation technology. The semipermeability of the porous morphology is based essentially on the spatial cross-section of the permeating species, ie, small molecules exhibit a higher permeability rate through the fiber wall. While the anisotropic morphology of the dense membrane which exhibit the dense skin, is obtained through the solution-diffusion mechanism. The permeation species chemically interacts with thepolymer matrix and selectively dissolves in it, resulting in diffusive mass transport along the chemical potential gradient, as what demonstrated in the pervaporation process.

Advantages and Disadvantages of Membranes

Hollow fiber is one of the most popular membranes used in industries. It is because of its several beneficial features that make it attractive for those industries. Among them are:{parse block="google_articles"}

• Modest energy requirement: In hollow fiber filtration process, no phase change is involed. Consequently, need no latent heat. This makes the hollow fiber membrane have the potential to replace some unit operation which consume heat, such as distillation or evaporation column.
• No waste products: Since the basic principal of hollow fiber is filtration, it does not create any waste from its operation except the unwanted component in the feed stream. This can help to decrease the cost of operation to handle the waste.
• Larger surface per unit volume: Hollow fiber has large membrane surface per module volume. Hence, the size of hollow fiber is smaller than other type of membrane but can give higher performance. 
• Flexibility: Hollow fiber is a flexible membrane, it can carry out the filtration by 2 ways, either is "inside-out" or "outside-in".
• Low operating costs: Hollow fiber need low operation cost compare to other types of unit operation.

However, it also have some disadvantages which lead to its application constraints. Among the disadvantages are:

• Membrane fouling: Membrane fouling of hollow fiber is more frequent than other membrane due to is configuration. Contaminated feed will increase the rate of membrane fouling, esapecially for hollow fiber.
• Expensive: Hollow fiber is more expensive than other membrane which available in market. It is because of its fabrication method and expense is higher than other membranes.
• Lack of research: Hollow fiber is a new tachnology and so far, research done on it is less compare to other types of membrane. Hence, more research will be done on it in future because of its potential.
• Cleaning and temperature restraints: Since hollow fibers of made of polymers there are constraints with regards to what types of chemicals they can contact and the temperatures at which they can operate.

Membrane Process

Various types of membrane processes can be found in almost all of the literature references. In this text, we will confine ourselves to the few membrane processes that we will encounter in the further discussion of the industrial applications.

Reverse Osmosis (RO)

Figure 1: Membrane Configurations

There is considerable confusion in the open literature as to the distinction between few membrane separation processes, i.e., the microfiltration (MF), ultrafiltration (UF) and reverse osmosis (RO). Occasionally one will see it referred to by other names such as "hiperfiltration (HF)". In order to distinguish these separation processes clearly, Porter in his paper presented one of the useful method based on the smallest particles or molecules which can be retrained by the various membranes. Accordingly, RO has the separation range of 0.0001 to 0.001mm (i.e., 1 to 10 Ã… ) or < 300 mol wt.

RO is a liquid-driven membrane process, with the RO membranes are capable of passing water whilst rejecting microsolutes, such as salts or low molecule weight organics (< 1000 daltons). Pressure driving force (1 to 10 MPa) needed to overcome the force of osmosis that cause the water to flow from dilute permeate to concentrated feed. The principle use of this membrane process is desalination, which show its great advantage over the conventional technique of desalination, i.e. ion exchange.

Pervaporation (PV)

In this process, liquid mixture are fed under pressure to a non-porous membrane, where components pass through the membrane by solution-diffusion and evaporate at the permeate side of the membrane. This technique is able to separate an azeotropic mixture. It current usage is well know in dehydration of the organic solvents and mixtures and the removal of organics from aqueous stream. The future application of this process, which is now under the main interest of the researcher is the hydrocarbon separation^,, which shows its advantages of energy require compared to the conventional distillation technique.

Gas Separation

Two type of gas separation processes have been encountered: gas permeation (GP) and gas diffusion (GD). The gas separation of the industrial interest is the former process, which is a pressure driven process where vapor components pass through a non-porous membrane by a solution-diffusion mechanism; analogous to RO. While gas diffusion process can be done for the microporous membranes, operating under a concentration or partial pressure gradient.

Chemical and Petrochemical Applications

Gas Separation

Gas membranes are now widely used in variety of application areas, as shown in Table 2. This is because of its advantages in separation, low capital cost, low energy consumption, ease of operation, cost effectiveness even at low gas volumes and good weight and space efficiency.

As the matter of fact, hollow fiber is playing a important role in gas separation. It is because of its high separation areas and selectivity. {parse block="google_articles"}The hollow fibers have approximately 30 times the productivity of other oxygen enriching membranes plus excellent inertness associated with their totally flourinated chemistry. The market of the gas separation include, small and intermediate scale industrial oxygen and nitrogen at moderate purity levels(oxygen 25%-40% or nitrogrn 82%-95%), portable oxygen for respiratory care, enhanced engine power and emissions reduction and removal gases from liquid.

Hollow fibers have demonstrated stable, high flux with moderate selectivity in full scale system. The high flux from hollow fibers is due to the combination of high transfer or separatin areas and thin membrane wall. Besides, it also has a low surface energy.

With such characters, hollow fiber is widely used in many gas separation industries. For instance, it is used in O₂/N₂ separation for oxygen enrichment and inert gas generation, H₂ /hydrocarbons separation for refinery hydrogen recovery, H₂ /CO separation for sygas ratio adjustment, H₂/N₂ separation for ammonia purge gas, CO₂/hydrocarbons separation for acid gas treatment and landfill gas upgrading, H₂O/hydrocarbons separation for natural gas dehydration, H₂S/hydrocarbons separation for sour gas treating, helium separation and etc.(Table 3).

Apart from that, low capital cost of hollow fiber also lead to its popularity. For example, for oxoalcohol feed separation, the process cost is about 1.000 for hollow fiber membrane. However, for crygenic(partial condensation) and PSA processes are about 1.234 and 1.133 respectively.(appen ) From here, we can see that most of the cost for hollow fiber is for compression and not for purification. It is because hollow fiber itself already provides a good medium for purification.

Desalination

As mentioned in the above section, RO is mainly use to remove the dissolved ion in the feed water. Its current extensive use in Malaysia industry sector is found in the production of ultrapure water in the semiconductor manufacturing industry. Historically, distillation and ion exchange was first used to remove the inorganic salts, but RO membrane processes with the combination of ion-exchange system has promised a better result in both the product requirement and a better economic view point.

Other usage of RO included the removal of organics, salts and silica ahead of deionizers in boiler feed water, removal of inorganic salts, phosphorus and nitrogen compound in the municipal waste water treatment and also the demineralization of sea water and brackish water in the production of potable water. Porter provide a good reference in the comparison of product quality and economic of the above processes.

Biotech and Biochemical Applications

The biotechnology industry, which originated in the late 1970s, has become one of the emerging industry that draws the attention of the world, especially with the emergence of the genetic engineering as a means of producing medically important proteins, during the 1980s. Two of the major interest applications of membrane technology in the biotechnology industry will be the separation & purification of the biochemical product, as often known as Downstream Processing; and the membrane bioreactor, which developed for the transformation of certain substrates by enzymes (i.e. biological catalysts).

Downstream Processing

{parse block="google_articles"}"Downstream Processing", a new key term given a decades ago, devotes towards the science and engineering principles in separation and purification in this emerging industry, has become a key issue to enhance the quality of the biochemical product. It is particularly important because it typically accounts for nearly three-fourths of the manufacturing costs in this new industry and because reliable and effective purification can be of the utmost important to the user. Membrane separation, together with the bioaffinity chromatography, liquid extraction and selective precipitation are the few techniques in the bio-separations, which gain attention from both the industries and researcher in order to upgrade the product quality of the biochemical industry.

Lots of study has been put in this area involving the most of the recovery of the biofuels and the biochemicals. Throughout the available literature, the most useful review is presented by Stephen A Leeper (1992), which compiles a large number of the previous and current studies' data on the different types of biofuels and the biochemicals product recovery, consisting of the usage of different types of membrane materials, membrane processes, together with the operating parameters of the studies being carried out.

Membrane Bioreactors

Since its introduction in the 1970s, membrane bioreactor has granted a lot of attention over the other conventional production processes is the possibility of a high enzyme density and hence high space-time yields. Whereas downstream processing is usually based on discontinuously operated microfiltration, membrane bioreactor are operated continuously and are equipped with UF membranes. Two type of bioreactor designs are possible: dissolved enzymes, (as in used with the production of L-alanine from pyrurate) or immobilized enzymes membrane.

Future Prospects

Membrane science began emerging as an independent technology only in the mid-1070s, and its engineering concepts still are being defined. Many developments that initially evolved from government-sponsored fundamental studies are now successfully gaining the interest of the industries as membrane separation has emerged as a feasible technology.{parse block="google_articles"}

As were noted by the US National Research Council, the technological frontiers of the membrane technology should be concerned more in the developing of new membrane materials and the identification of new ways of using permselective membranes.

New membrane materials to be used is still a big option in the research of this brand new technology, as most of the researchers are always intend to get a better improvement for this separation process. Journal of Membrane Science serve as a good reference, where lots of the new membrane materials research may be found.

For the latter, membrane-based hybrid system serve as a good example, as it is a combination of conventional unit operations and membrane separation processes, which often results in separation processes that offer significant advantages over the exclusive use of either component process. Such advantages may include more complete separation, reduced energy requirement, lower capital cost, and lower production cost. Two good example of this hybrid system are the RO / evaporator hybrid system to concentrate corn steep water, and a membrane / vapor-recompression hybrid process to recover energy in hot, moist dryer exhaust. Studies has also been carried out and proven that these hybrid system did perform a better of the washwater purification and reuse pilot plant (HUMEF) which has been successful installed in Eindhoven pumping station, the Netherlands.

References

1. Parker, Sybil. P, 1994, McGraw Hill Dictionary of Scientific and Technical Terms, McGraw Hill.
2. Philip A. Schweitzer, 1988, Handbook of Separation Technique for Chemical Engineers, 2^nd Edition, McGraw Hill.
3. Douglas M. Ruthven, 1997, Encyclopedia of Separation Technology, John Wiley & Sons.
4. Jacqueline I. Kroschwitz, 1991, Concise: Encyclopedia of Polymer Science and Engineering, John Wiley & Sons.
5. Howley, Gessner G., 1977, The Condensed Chemical Dictionary, 9^th Edition, Van Nostrand Reinhold Company. , 7^th Edition, Volume IX, Clarendon Press Oxford, 1989.
6. The Oxford English Dictionary
7. Green, Perry, 1997, Perry’s Chemical Engineers’ Handbook, 7^th Edition, McGraw Hill,. [Section 22-37 – 22-69]
8. National Research Council (US), 1989, Frontiers in Chemical Engineering: Research Needs and Opportunities,.Washington DC: National Avademy.
9. Cabasso, Gulf South Israel Research Institute.
10. Fauzi I., Ghazali N., Rosli Y., 1998, A Short Course of Membrane Technology, CEPP, University Technology Malaysia.
11. R. Rautenbach, R. Albrecht, 1989, Membrane Processes, John Wiley & Sons. , Hungarian Chemical Society. 1996, University of Indonesia.
12. Proceeding for the 7^th World Filtration Congress Budapest, Hungary 1996
13. Proceeding of the Regional Symposium of Chemical Engineer,
14. R. E. Kesting / A. K. Fritzsche, 1993, Polymeric Gas Separation Membranes, John Wiley & Sons.
15. Ralph E. W., Peter N. P. (Eds), 1986, Industrial Membrane Processes, AIChE Symposium Series (No. 248, Vol82), AIChE.
16. Peter A. (Ed), 1998, World Water and Environmental Engineering, Vol.21, Issue 3, March 1998.
17. Kamalesh K. S., Douglas R. L., 1988, New Membrane Materials and Processes for Separation, AIChE. IChemE Symposium Series, 1978, IChemE.
18. Alternatives to Distillation,
19. Reynolds, Richard, 1996, Unit Operations and Processes in Environmental Engineering, PWS Publishing Company.
20. Hamdani S., Membrane Technology: The Right Choice for ASEAN, UTM.
21. Roger G. H. (Ed), 1994, Protein Purification Process Engineering, Marcel Dekker.
22. Gomez-Fernandez J., D. Chapman, L. Packer, 1991, Progress in Membrane Biotechnology, Birkhauser.
23. David S. Soane (Ed), 1992, Polymer Applications for Biotechnology: Macromolecular Separation and Identification, Prentice-Hall.
24. G. Street, 1994, Highly Selective Separations in Biotechnology, Blackie Academic & Professional.
25. Jean-Francois. H., Jean H., Subhas S., 1990, Downstream Processing and Bioseparation: Recovery and Purifucation of Biological Product, American Chemical Society.
26. Norman L., Joseph C., 1992, Separation and Purification Technology, Marcel Dekker.
27. M.S. Verrall and M. J. Hudson, 1987, Separations for Biotechnology, Ellis Horwood Limited.
28. Munir Cheryan, 1989, Ultrafiltration in Food and Bioprocessing, University of Illinois.
29. W. E. L. Spiess, H. Schubert, Engineering and Food - Advanced Processes, Vol. 3, Elsevier Applied Science.
30. J. Krijgsman, 1992, Product Recovery in Bioprocess Technology, Butterworth-Heinemann.
31. J. D. Stowell, P. J. Bailey, D. J. Winstonley (Eds.), 1986, Bioactive Microbial Product 3: Downstream Processing, Academic Press.
32. J. P. Hamel, Jean B. Hunter, Subhas K. Sikdar (Eds.), 1990, Downstream Processing and Bioseparations: Recovery and Purification of Biological Products, American Chemical Society.
33. http://www.aces.uiuc.edu/~fshn/faculty/cheryan.html [under Research]
34. http://www.dupont.com/
    

There are many different types of equipment used for leaching.  Most of these pieces of equipment fall into one of two categories:{parse block="google_articles"}

Percolation ("Liquid added to solids") - The solvent is contacted with the solid in a continuous or batch method.  This method is popular for in-place ore leaching or large scale "heap" leaching.  Popular for extreme amounts of solids.

Dispersed Solids ("Solids added to liquid") - The solids are usually crushed into small pieces before being contacted with solvents.  This is a popular leaching method when an especially high recovery rate can economically justify the typically higher operating cost (Ex/ gold extraction)

Whether the leaching is taking place via percolation or by dispersed-solids, there are three important factors that aid in leaching:   temperature, contact time/area, and solvent selection.  Temperature is adjusted to optimize solubility and mass transfer.

Liquid-to-solid contact is essential for the extraction to take place and maximize contact area per unit volume reduces equipment size.   Solvent selection plays an important role in solubilities as well as the separation steps that follow leaching.  Nearly all leaching equipment employs some type of agitation to aid in mass transfer and to ensure proper mixing.  The most popular leaching equipment can be seen in Perry's Chemical Engineers' Handbook.

General Arrangement and Nomenclature

The following nomenclature is used in conjunction with Figure 1 below:

S_v = Mass flow of solvent
S_t = Mass flow of solute
O_flow = Mass flow of solvent + mass flow of solute (Overflow)
F_ins = Mass flow of insoluble solids
F_sol = Mass flow of soluble solids + residual solvent (Underflow)
N = F_ins / F_sol
T = Total solution flowrate
X_i = Mass fraction of solute in O_flow at a given stage
Y_i = Mass fraction of solute in F_sol at a given stage

Figure 1: Four Stage, Countercurrent Leaching

Material Balance

Let's begin with a mass balance on the solute or the material that is being removed via leaching:

Eq. (1)

Solving for X o yields the operating line equation:

Eq. (2)

 Next, we perform an insoluble solids balance:

Eq. (3)

At this point, we introduce the graph construction.  A plot of N vs. X,Y is used to step off stages for leaching calculations.  But, just as equilibrium data is necessary for a McCabe-Thiele diagram in distillation, leaching calculations require that you know something about how the solids and liquids interact.   Settling experiments can provide data such as those shown in Table 1.  Essentially, N = Fins / Fsol so as more solvent is mixed with the solids, Fsol increases, N decreases, and Y increases.  A plot of this data may resemble Figure 2.

Figure 2: Sample Plot of Experimental Data

Once this data is obtained, four distinct points are known and can be plotted:  Fsol o, Oflow 1, Oflow o, and T.  For example, let's say that 1000 kg/h of solids, wetted with 100 kg/h of solvent, will be fed to a leaching system and of this amount 400 kg/h are soluble in the solvent.  The 1500 kg/h of lean solvent coming from the separation section contains 5 wt % solute.  The desired mass fraction of solute leaving the leaching system is 0.55.  All of these values are determined by systems outside the leaching equipment or they are dependent on the leaching solvent, operating temperature, or particle size.  For the example above, our graph would resemble Figure 3.

Figure 3: Finding Fsol at the End of the Leaching Process

Point T is found by:

We still have not addressed how to find the total number of stages required and intermediate solute concentrations.  In order to do this, we introduce the operating point equation.

P = operating point or difference in flows
Np = N-coordinate of the operating point
Xp = X-coordinate of the operating point

Eq. (4)

Eq. (5)

Eq. (6)

The operating point is now used to construct tie lines for the intermediate stages as shown in Figure 4.

Figure 4: Leaching Operating Graph for a Four Stage Leaching Operation

Parting Word

Leaching calculations are at times confusing due to the "clumsy" nomenclature and the physical substances involved.  Once you've identified the variables and the experimental data that you have and you're able to construct a graph such as Figure 3 the remaining steps are relatively simple.  When constructing Figure 4, be sure that the graph scale is sufficiently large to plot the operating point, it's your guide to the remainder of the process.

Realize that to optimize a leaching process you may have to evaluate many solvents, particle sizes, operating temperatures, and feed compositions.

The first step in designing a packed tower is more science than art. {parse block="google_articles"}The equilibrium data between the contaminant and the solvent (or the distillation components) is needed for the analysis. If tabulated data for your system is unavailable and the total amount of the contaminant is small (as it usually will be), Raoult's Law can be used to estimate the equilibrium data for absorption or stripping applications. For distillation, equilibrium data can be predicted by selecting the appropriate thermodynamic model. The operating line for the tower is constructed differently depending on whether you're dealing with distillation or absorption/stripping.

Figure 1: General Packed Column Arrangement

Since we're focusing on absorption, we'll use it as an example. In absorption/stripping, the operating line is constructed differently depending on whether the contaminated stream can be considered "dilute" or if it must be treated as a concentrated stream. Usually, it is safe to treat the stream as dilute if the contaminant makes up less than 10 mole percent of the stream. For streams that cannot be considered dilute, the mass transfer coefficients must be evaluated in terms of the gas and liquid flows. Then, graphical evaluation of several integral relationships must be completed. This type of evaluation is outside the scope of this article and a text should be consulted for solving these types of problems. For this article, we will consider dilute streams which are more common for packed tower absorption and stripping.

Dilute streams allow the column designer to assume constant mass transfer and the operating line can be constructed in terms of the simplified balance shown below:

Quick Example

Suppose you wish to remove acetone from a gas stream of 10,000 mol/h in a packed column. The inlet gas contains 2.6 mole percent acetone and the outlet gas stream can contain no more than 0.5 mole percent acetone. Assume a pure water stream enters the packed tower at a rate of 8,000 mol/h.

Figure 2: Example Packed Column

L out x out + G out y out = L in x in + G in y in
(8000) x out + (10000)(0.005) = (8000)(0)+(10000)(0.026)

x out = 0.02625

The equilibrium and operating lines are constructed as follows:

Just as in the McCabe-Thiele analysis of distillation, the equilibrium stages are stepped off between the two lines. Note that for stripping, the operating line would be on the other side of the equilibrium line.

Once the theoretical number of stages have been determined, you can proceed with the design of the column by following the three steps that we'll outline below.

Figure 3: Constructing the Operating and Equilibrium Lines

Detailed Step-by-Step Example

Specify the packing type and column dimensions for a column that will be used to remove chlorine from a gas stream using an organic solvent. Assume the separation requires 20 theoretical stages. The vapor flow is 7000 kg/h, the average vapor density is 4.8 kg/m³. The liquid flow is 5000 kg/h, the average liquid density is 833 kg/m³. The liquid's kinematic viscosity is 0.48 centistokes (4.8 x 10^-7 m²/s) {parse block="google_articles"}

Step 1: Selecting the Type and Size Packing

This is where the art of designing packed columns begins. Some people believe that there are stringent rules surrounding the choice between random and structured packing. You can think of random packing as the type that comes in a sack and it is simply dumped into the column. Structured packing may come in bales or intricate designs that are stacked in specific patterns. This is probably one of those areas of engineering where past experience in the application is the best guide. Two "areas of choice" where structured packing is used are in very low pressure drop applications and for increasing the capacity of an existing column. Since we're considering a new design with no serious pressure drop constraint, we'll choose the more economical random packing.

Table 1 shows both English and Metric unit packing factors for some common packings.

Generally, the column diameter to packing size ratio should be greater than 30 for Raschig rings, 15 for ceramic saddles, and 10 for rings or plastic saddles. The geometry of your packing will typically be a function of the needed surface area and/or allowable pressure drop. If several packings meet your requirements, you'll typically choose the least expensive so long as it has an acceptable operating life. For our example, we'll choose Pall rings (plastic). For columns over 24 inches in diameter, No. 2 or 2 inch packing should be examined first. By looking at our flowrates, the chances of our column having a diameter of at least 24 inches are good, but we'll verify this later. For now, we'll settle on 2 inch plastic Pall rings for our initial analysis.

Step 2: Determine the Column Diameter

Most methods for determining the size of randomly packed towers are derived from the Sherwood correlation. A design gas rate, G, can be determined with the help of the figure below which is based on correlation from the Sherwood equation.

Each line on the graph is marked with an acceptable pressure drop in inches of water per foot of packing (numbers in parentheses are in mm of water per meter of packing).

Figure 4: Design Gas Rate Chart

Guidelines are as follows:

• Moderate to high pressure distillation = 0.4 to 0.75 in water / ft packing
  = 32 to 63 mm water / m packing

• Vacuum Distillation = 0.1 to 0.2 in water / ft packing
  = 8 to 16 mm water / m packing

• Absorbers and Strippers = 0.2 to 0.6 in water / ft packing
  = 16 to 48 mm water / m packing

These guidelines are designed around "flooding pressure drops" documented in literature. In other words, for most cases, designing with these pressure drops should help you avoid flooding. In the later stages of design, you may want to perform a thorough flooding calculation. Perry's Chemical Engineers' Handbook covers this topic well. Since we are designing an absorber, we will design for 42 mm water / m packing (you could design for a lower pressure drop, but the column will increase in diameter and most likely cost). First, we'll evaluate the x-axis of the graph above:

Figure 5: Reading the Design Gas Rate Chart

(L/V)(vapor density/liquid density)^0.5 = (5000/7000)(4.2/833)^0.5 = 0.0507

Note that 4.2 kg/m³ was used for the vapor density. The average vapor density was given as 4.8 kg/m³. However, at the top of the column, the vapor will be less dense and at it's highest velocity. This is what you should design for. As a rule of thumb, I reduce the average vapor density by about 15% for design, however if you can get real data from a similar tower, certainly do so! Reading the intersection of the 42 mm water/m packing line and 0.05 on the axis, we find a value of 1.5 for the y-axis (see Figure 5).

From the previous charts, we read a packing factor of 24 for 2 inch plastic Pall rings. All other information is know so we can solve for G as shown on the y-axis of the graph:

G = [1.5 [(4.2)(833-4.2)]/[(10.764)(24)(0.48)^0.1]]^0.5 = 4.66 kg/m² s

Now, we solve for the column cross sectional area:

Ax = Vapor Flow / G = 7000 kg/h / [(4.66 kg/m² s)(3600 s/hour)] = 0.42 m²

and the column diameter is calculated by:

Diameter = [Ax / (PI/4)]^0.5 = [0.42/(PI/4)]^0.5 = 0.73 m or 2.4 ft

So our assumption of at least a 24 in column diameter is accurate. If it had not been accurate, G would be recalculated using a smaller packing which would also correspond to a larger packing factor.

Step 3: Determine the Column Height

Perhaps the most interesting step in designing a packed column is deciding how tall to build it. You should first ask yourself "What stage of the design are we currently working on?" If the design is preliminary, the general HETP (Height Equivalent to a Theoretical Plate) will work well. If the design requires a higher degree of accuracy, I recommend consulting the packing manufacturer or a book entitled Distillation Design by Henry Kister (McGraw-Hill, ISBN 0-07-034909-6). Distillation Design contains an exhaustive list of HETP values based on the components of the system and the type of packing used (Chapters 10 and 11). As for preliminary estimates, the following HETP values should be used:

To determine the height of the absorption tower in our example, we multiple the 20 theoretical stages by 6 ft or 1.83 m. We estimate the height of the tower to be 120 ft or about 37 meters.

Other Notes

While our example problem focused on absorption, packed towers are also widely used in distillation. Perhaps the most popular of which is the well documented vacuum distillation of ethylbenzene and styrene in the production of styrene. Distillation Design covers this application very well.

Pervaporation can used for breaking azeotropes, dehydration of solvents and other volatile organics, organic/organic separations such as ethanol or methanol removal, and wastewater purification.

Characteristics of the pervaporation process include:

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Figure 1: Overview of the Pervaporation Process

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Figure 2: Schematic of Liquid Permeation

1. Low energy consumption

2. No entrainer required, no contamination

3. Permeate must be volatile at operating conditions

4. Functions independent of vapor/liquid equilibrium

Types of Pervaporation Processes

Batch pervaporation is a simple system with great flexibility, however a buffer tank is required for batch operation.  Continuous pervaporation consumes very little energy, operates best with low impurities in the feed, and is best for larger capacities.   Vapor phase permeation is preferred for direct feeds from distillation columns or for streams with dissolved solids.

Pervaporation for Separation

Liquid transport in pervaporation is described by various solution-diffusion models¹.   The steps included are the sorption of the permeate at the interface of the

solution feed and the membrane, diffusion across the membrane due to concentration gradients (rate determining steps), and finally desorption into a vapor phase at the permeate side of the membrane.  The first two steps are primarily responsible for the permselectivity¹.  As material passes through the membrane a "swelling" effect makes the membrane more permeable, but less selective, until a point of unacceptable selectivity is reached and the membrane must be regenerated. 

The other driving force for separation is the difference in partial pressures across the membrane.  By reducing the pressure on the permeate side of the membrane, a driving force is created.  Another method of inducing a partial pressure gradient is to sweep an inert gas over the permeate side of the membrane.  These methods are described as vacuum and sweep gas pervaporation respectively.

Basics of the Pervaporation System

Figure 3 shows a typical pervaporation system.  The feed is allowed to flow along one side of the membrane and a fraction of the feed (permeate) passes through the membrane and leaves in the vapor phase on the opposite side of the membrane.   The "vapor phase" side of the membrane is either kept under a vacuum or it is purged with a stream of inert carrier gas.  {parse block="google_articles"}The permeate is finally collected in the liquid state after condensation.  The liquid product is rich in the more rapidly permeating component of feed mixture.  The retentate is made up of the feed materials that cannot pass through the membrane.

Figure 3: Simplified Pervaporation  Process

Membranes

The membranes used in pervaporation processes are classified according to the nature of the separation being performed.  Hydrophilic membranes are used to remove water from organic solutions.  These types of membranes are typical made of polymers with glass transition temperatures above room temperatures.  Polyvinyl alcohol is an example of a hydrophilic membrane material.  Organophilic membranes are used to recover organics from solutions.  These membranes are typically made up of elastomer materials (polymers with glass transition temperatures below room temperature).   The flexible nature of these polymers make them ideal for allowing organic to pass through.  Examples include nitrile, butadiene rubber, and styrene butadiene rubber.

Factors Affecting Membrane Performance

According to the solution-diffusion model, higher fluxes can be obtained with an increased thermal motion of the polymer chains and the diffusing species.  Properties of the polymers that affect diffusion include the "backbone" material, degree of cross-linking, and porosity.  Molecular-level interactions between membranes and diffusing species is expressed via a permeability constant used in the Arrhenius relationship:

Eq. (1)

Where,
E_p = Activation energy
P_o = Permeability constant
R  = Gas constant
T  = Temperature

Pervaporation Characteristics

Molecular Flux

Molecular flux is the amount of a component permeated per unit area per unit time for a given membrane.

Eq. (2)

Where,
J_i = Flux of component "i" (moles/h cm²)
Q_i = Moles of component "i" permeated in time "t"
A = Effective membrane surface area (cm²)

Permselectivity

The performance of a given membrane can be expressed in terms of a parameter called permselectivity:

Eq. (3)

Eq. (4)

Assuming the density of the components in the feed is the same, then:

Eq. (5)

Where,
X = Weight fraction
V = Volume fraction
p = Density
Superscripts "p" and "f" denote "permeate" and "feed" respectively while "i" and "j" represent individual components.

Permeability Coefficient

The molecular flux for pervaporation across a membrane can be related to the permeability coefficient by:

Eq. (6)

or

Eq. (7)

Here and , therefore

Eq. (8)

Equation 6 becomes:

Eq. (9)

Eq. (10)

where:

Established industrial applications of pervaporation include:
The treatment of wastewater contaminated with organics⁴
Pollution control applications⁴
Recovery of valuable organic compounds from process side streams⁵
Separation of 99.5% pure ethanol-water solutions⁶
Harvesting of organic substances from fermented broth⁷

Other products separated or purified by pervaporation include:

Continuing Research on Pervaporation

Pervaporation of Apple Juice

Pervaporation is used to recover any lost juice solution during evaporation.  The vapor from the evaporation process is further processed using pervaporation.  The recovered, concentrated apple juice can be combined with the product solution to help the apple juice retain it's aromatic and taste qualities.

Pervaporation in the Production of Fuel Ethanol

To establish a continuous fermentation process, the ethanol concentration within the fermentation vessel must be kept at 5% by weight or lower.  Pervaporation has been used to maintain the necessary ethanol concentration in the broth.  The advantages of using pervaporation in such a system include the ease of processing the clean, nearly pure ethanol extracted from the fermentation vessel and a significantly higher fermentation capacity or the reduction in fermentor size and costs.

Summary

Pervaporation continues to evolve as a feasible separation technology for many different applications.  As a proven method of separation as low temperatures and pressure, further application development for food processing is likely.  Using pervaporation to clean wastewater streams by removing a variety of organic compounds also holds much promise.

References

1.    Yong Soo Kang, Sang Wook Lee, Un Young Kim and Jyong sup shim, Pervaporation of water – Ethanol mixtures through
       cross – linked and surface modified poly (vinyl alcohol) membrane, J. Member. Sc., Elsevier Science Publishers B.V., Amsterdam, 51, 215, 1990.
2.    K.W. Boddeker and G. Bengston, Pervaporation membranes separation processes, Ed. By R.Y M. Hang. Elsevier, Amsterdam 437 – 460, 1991.
3.    G.H. Koops and C.A. Smolders Pervaporation membrane separation process, Ed. by R.Y.M Haung, Elsevier, Amsterdam 249 –273, 1991.
4.    C. Lipski and P. cote, the use of Pervaporation for removal of organic containment from water, Environmental program, 9, 254 –261, 1990.
5.    J. Kashemekat, J.G. Wiljmans and R.W Baker, Removal of organic solvent containments from industrial effluent streams by        
       Pervaporation, Ed. By R. Bakish, Proc. 4th int. Conf. On Pervaporation, process in chemical industry, Bakish materials    
       Corporation, Englewood, NJ, 321, 1981.
6.    B.K. Dutta and S.K Sridhar, separation of azeotropic organic liquid mixtures by Pervaporation, AIChE journal, vol.37, No.4, 581– 588, 1991.
7.    M.E.F. Garcia, A.C. Habert, R. Nobrega and L.A. Piers, Use of PDMS and EVA membranes to remove ethanol during      
       fermentation, Ed, by R. Bakish Proc. 5th Int. Conf. on Pervaporation process in the chemical industry, Bakish Materials      
       corporation, Englewood, NJ, 319 – 330, 1991.
8.    Aptel, P., N. Challard, J. Cuny, and J. Neel, “ Application of the Pervaporation Process to Separate Azeotropic Mixtures,” J.   
       Membrane Science., 1, 271 (1976).
9.    Dutta, B.K., D. Randolph, and S.K. Sikdar, “ Separation of Amino Acids Using Composite Ion Exchange Membranes,”   
       Biochemical Engineering VI, new York Academy of science, 589, 203,1990

Figure 1: Distillation Column

We'll examine this column and with some help from the Tower Pricing and Sizing Add-In, we'll see how small changes can save you money in energy costs.

Check the Product Purity

Rating: Little to no investment
Many companies tend to excessively purify products and sometimes with good reason. However, purifying to 98% when 95% is acceptable is just not necessary. In this case, the reflux rate should be decreased in small increments until the desired purity is obtained. In our example, the current reboiler duty will require $68,000 in low pressure steam per year. However, if the purity is decreased and the duty in the reboiler decreases by 5% (down to 2545 MJ/h), the cost of steam decreases to $64,500.

Summary: Excess purification was causing increased spending in the amount of $3,500 per year in the reboiler alone.

Seasonal Operating Pressure Adjustment

Rating: Little or no investment
For plants that are in locations that experience winter climates, the operating pressure can be reduced according to a decrease in cooling water temperatures. Although lowering the pressure in our column would actually be an expense since it's under vacuum, this would not always be the case. The lower pressure will facilitate separation thus lowering energy consumption.

Summary: Take advantage of lower temperatures outside.

Great attention to detail is required in the design of packed towers to achieve the necessary absorption efficiency. In many ways, the sulfuric acid industry is unique in that packed towers of exceptionally large{parse block="google_articles"} diameters with relatively small packing height are common. In addition, the use of large size ceramic packing has become the industry standard. A result of the unusual features of packed towers employed in acid plant service is that truly applicable design data are not readily available and that discrepancies reveal themselves when the designs of different technology suppliers are compared. For the engineer facing the task of sizing a packed tower or selecting a supplier, it is tempting to assume that the design techniques for acid plant towers are well proven and that differences in supplier''s offers simply reflect differences is design conservatism.

It is suggested, however, to take into consideration the following points:

A standard three inch saddle is available from a number of suppliers at relatively low costs. This saddle has been used for over thirty years in the sulfuric acid industry. Only modest profit margins in making and supplying this type of packing can be expected, not sufficient to commission significant development work on packing performance in large towers, especially when competing suppliers would gain the benefit of the development at no cost. Contractors are in a similar position, leaving the owners as the only party likely to gain from development work. Where new packing is proposed, there is a need to compare it with existing packing to see if there is an improvement which justifies the development expense or makes the changeout of packing attractive to the owners. This economic reality has limited the introduction of new packing over the past twenty years.

Engineers involved in sizing towers have at their disposal a number of different techniques for tower sizing, ranging from rules of thumb based on gas velocity and irrigation rate, to dated theoretical work in the

Figure 1: A Typical Sulfuric Acid Plant Layout (Click to Enlarge)

handbooks, to software programs from packing vendors, and finally to proprietary in-house design techniques. The resulting tower sizes vary significantly, as this paper will show. In addition, there is a need for design approaches which can be used with new packing for which there is little data. Most of the experimental work on packing pressure drop was carried out over forty years ago, almost exclusively in small pilot towers. Norton, for example, did much of their early work using thirty inch diameter towers while Koch used a thirty-six inch diameter tower. When the packing sizes were relatively small, the effect of the tower diameter on the packing density was minimal, but when larger packings were used in these pilot towers, there were significant edge effects and the void fraction in the test column was much larger than that found in large towers typical in acid plants with the same packing. The result was very optimistic predictions of pressure drop. Figure 2,

reproduced from a brochure published by VFF Industries, shows the relative number of pieces of packing per unit volume, the packing density, as a function of the ratio of the tower diameter to the characteristic dimension of the packing. The curve uses a reference packing density in a tower with a diameter twenty times the nominal size of the packing. For a three inch nominal size saddle, this would give a tower diameter of six feet. For a three foot tower, the packing density would be ninety seven percent of that of the reference tower, while for a twenty foot tower, the actual packing density would approach one hundred and ten percent of the reference case. The packing void fraction will vary accordingly with high packing densities resulting in low void fractions. The packed tower pressure drop and flooding limits are very sensitive to the void fraction, as will be shown later in this paper. It has been found that the pressure drop for a given large size packing in a plant scale tower can exceed twice the pressure drop measured in a pilot tower under identical process conditions.

Figure 2: Packing Density Correction Factor (Click to Enlarge)

The Generalized Pressure Drop Correlation (GPDC) is the classic sizing method for packed towers and is used in many industries. It is, however, based mostly on the small pilot tower data. As long as the correlation is applied to small packing, it appears to give reasonable results, but when the performance of large packing in large towers is assessed, then the results appear overly optimistic. A second rule of thumb method is to size the tower based on a packing exit gas velocity of 8 ft/s with an acid irrigation rate of 10 USGPM/ft². Most often the velocity used in sizing the tower is that of the gas leaving the packing. Different techniques should be applied to the tower bottom depending on the temperature of the inlet gas and the relative acid flow. A third sizing approach is to use the published pressure drop curves for the air and water system which are publicized in the suppliers'' literature. Figure 3 shows one such plot published by U.S. Stoneware. This approach is more practical but, as previously discussed, the data were again mostly developed in small pilot towers, even when large packing was studied. On the basis of proprietary design approaches, one from a packing supplier, and the second from a chemical company, a design approach was developed by CECEBE based on ideas originally proposed by Dr. Max Leva. Field data were carefully taken in full sized towers to refine the correlations and design methods. The method of Dr.

Figure 3: Pressure Drop versus Gas Rates 1 1/2" Intalox Saddles Ceramic (Click to Enlarge)

Leva starts with dry bed pressure drop, which is then corrected for the liquid flow through the packing and also for loading. With moderate gas flows and pressure losses, the incremental increase in pressure drop due to liquid irrigation depends on the liquid flow but does not change with gas velocity. Greater liquid flow causes greater liquid hold-up and higher interstitial gas velocities. These two effects combine to result in higher pressure drop. In this range of moderate gas and liquid flow rates, pressure drop curves for irrigated packing run parallel to the dry gas pressure drop curve. Further increases in gas and liquid flow, beyond a critical limit called "loading point", result in a rapid increase in the liquid hold-up due to spray re-entrainment and causes increased pressure loss. Eventually, the liquid hold-up fills a significant part of the packing void and it becomes difficult to force gas flow through the packing. The tower ultimately floods.

The various design approaches described earlier have been used in this paper in a standard absorber to predict the tower size on the basis of operating with standard packing. In addition, tower sizes were developed for both the HP^TM Saddle Packing developed by CECEBE Technologies and the Flexeramic® Structured Packing developed by Koch.

Packing Descriptions

The highly corrosive nature of sulfuric acid has restricted the materials that can be used economically in acid plants to ceramics. In the pioneering days of the industry, quartz rock or crushed brick were often used in packed towers. In due course these packings, which are very inefficient, were replaced by ceramic Raschig rings and by grid tile. The author has had personal experience with a situation where one tower with quartz rock was still found to be in operation while the others were filled with three inch cross-partition rings. Ceramic Pall Rings were also used in early days, but soon ceramic saddles became the industry {parse block="google_articles"} standard. Originally 1.5" saddles were used, followed later by 2" saddles, and finally by 3" saddles. Recently, CECEBE introduced the HP^TM Saddle Packing, which has more open structure and, therefore, has a much lower pressure drop with greater capacity. It is made of much stronger porcelain which has less tendency to form chips. Structured Ceramic Packing and Wave Packing have been introduced by other suppliers and design techniques need to be disclosed for these new packings. Surprisingly, the actual performance of the standard 3" saddle in a large tower also remains to be firmed up.

Physical Factors

 

Void Fraction

Since gas in an acid plant packed tower flows up and the liquid flows down to achieve the required gas-liquid contact, the packing functions as a liquid surface generation device. While good mixing action by the packing is a desired feature in that it promotes mass transfer, packing must not cause the tower to fill with liquid and prevent gas counterflow. The void fraction of packing, which is the space available for the gas and liquid to flow, depends on many factors. These factors include the shape of the packing, the diameter of the tower relative to the size of the packing, the degree of packing chips or sulfate accumulation and the liquid flow rate. Figure 2 illustrates the variation of packing density with the ratio of the tower diameter to packing characteristic dimension. Ceramic packings typically have void fractions around 0.75, although a number of references from small sized towers have declared 0.80 void fraction for the same packing. This small difference is significant in that a 0.05 decrease in void fraction will add over fifty percent to the pressure drop. Fouling with sulfate or chips can easily double the pressure drop as well and it may be desirable at the initial design stage to make an allowance for such fouling.

Effective Surface

While the actual surface of the packing pieces can be deduced from geometry, tests with Raschig rings have demonstrated that a significant portion of the surface is not wetted by the circulating liquid and is, therefore, ineffective. The same conclusion is not necessarily true for other packings. The HP^TM saddle and Pall Ring both have less surface than their predecessors but their surface is much more effective. With an identical shape, the surface required for satisfactory mass transfer is similar for small or large packing. The smaller packings, however, typically have a much larger surface area per unit packing volume and, therefore, will need less packing height. The penalty is that smaller packing has a larger pressure drop so that the tower diameter must be increased to handle a given gas and liquid flow rate. Note that liquid distribution also becomes more difficult as the diameter of a tower increases. In addition, tower costs are mostly a function of diameter and not of height. It is therefore, not economical to use small packing.

Gas Passage Size

For a given packing height, the number of time the gas must detour around a packing piece depends on the size of the packing, while the mass transfer depends to a significant degree on the extent to which the gas stream is split and contacts the liquid. The larger the packing, the higher is the height required for the mass transfer duty and the smaller the required tower diameter. The pressure drop depends largely on the number of gas detours as the gas passes through the packing.

Random Orientation

Packing can be structured or random. Structured packing is a relatively new development. It still remains expensive and difficult to install. Its capacity is, however, marginally higher than the standard 3" saddle. A serious shortcoming of structured packing is that good initial liquid distribution is absolutely necessary to achieve the required mass transfer efficiency. In other applications involving structured packing, ten irrigation points per square foot are commonly specified. In acid tower applications, a layer of saddle packing is often used on top of the structured packing to achieve adequate liquid distribution. This requirement voids the potentially higher capacity of structured packing. Most packings used in the sulfuric acid industry are random and the standard 3" saddle is a good example. The Berl Saddle, which preceded it, can stack in a tower and is now rarely used for that reason. The Raschig ring and derivatives are also good random shapes. For effective use, a packing dimension should be small compared to the tower in which it is installed and preferably packing should not have a long and thin shape. Such a geometry can lead to gas bypassing at the wall of the tower.

Mechanical Strength

Ceramic packing can break easily if it is not properly fabricated. The two techniques by which such packing is made are by extrusion and by slip casting. Extrusion is the most common technique as it is relatively inexpensive. Slip casting produces a saddle of greater density, which has much lower rates of absorption of acid and water and which is much harder to break. In addition, the packing shape must be designed to provide sufficiently thick cross sections to minimize breakage. The CECEBE HP^TM saddle benefits from both thick sections and the slip casting technique to give a packing that is exceptionally strong, nearly three times stronger than conventional ceramic 3" saddles. Structured packing uses thin sheets which are fragile.

A further factor affecting strength in general is the clay that is used in the manufacture. Most packing is made from mined clay with little further preparation and the clay composition can very depending on where in the pit it comes from. Clay used for porcelain or domestic fixtures is normally formulated to a specified grain size distribution and chemical composition and results in a stronger and more homogenous saddle. The HP^TM saddle, for example, uses the same clay that is used in the fabrication of toilets and sinks. It is of a much higher quality than the clay mined directly from the pit. This is important to an owner because packing can be broken during installation or during operation, resulting in chip generation which will then lead to excess pump wear, plugging of acid coolers and distributors, higher pressure drop, and expensive downtime for cleaning.

Acid Absorption

Where ceramics have been produced by extrusion, voids in the packing must be expected. This results in significant absorption of acid into the packing and extended acid weeping on shutdown, making access and maintenance difficult. Voids in slip cast ceramics are much less common. This difference reveals itself in the specific gravity of packing which can range from 2.3 in the extruded product to 2.7 in the slip cast porcelain.

Pressure Drop Predicition

When packed towers were initially introduced as mass transfer devices, the dry gas pressure drop across them was the first criterion which was evaluated and work by Ergun and Leva resulted in a correlation for dry bed pressure drop of the following form:

 

Eq. (1)

A detailed definition of Equation 1 can be found in Perry''s Chemical Engineers'' Handbook, 6th Edition. For packed towers using large size packing, the flow is generally turbulent, which results in "n" being equal to 2. Equation 1 can thus be simplified to:

Eq. (2)

C₂ combines the constants of Equation 1 and the characteristic packing dimension D_p.

Tower packings typically have a void fraction of 0.6 to 0.8, which makes the void fraction term in the denominator very significant. The gas velocity G can be replaced by the empty tower approach velocity, V. The term V/gives the interstitial gas velocity in the tower V_a so that Equation 2 can be re-written as:

Eq. (3)

Introducing a characteristic term Fs, where

Eq. (4)

results in

Eq. (5)

Figure 4 shows dry bed pressure drop curves for a variety of packings. These curves are based on both plant measurements and published data, but differ from handbook information in that the data have been adjusted to apply to full sized towers. The correction to the data has resulted in significantly higher pressure losses for the larger

Figure 4: Pressure Drop vs. Gas Velocity for Industrial Towers (Click to Enlarge)

packings. It is seen that each packing has an associated constant to express the relationship between the pressure drop and the gas velocity.

As already mentioned and shown in Figure 3, irrigating the packing with liquid results in liquid filling part of the void space and constricting the passages for gas flow. Leva introduced an exponential term for this phenomenon from experimental pressure drop data and CECEBE has continued with this approach and has further refined the equations based on full scale plant measurements. The equation resulting from this evaluation for the liquid correction is the following:

 

 

where

LC = "Liquid Correction" factor associated with irrigation below the loading point
V_L = irrigation rate (ft/s) based on an empty tower
C₃ = packing specific constant

The irrigated packing pressure drop is then,

 

Eq. (7)

Values of C₂ and C₃ for various packings were developed by correcting published pilot tower data for the smaller void fractions expected in full sized towers. The results are shown in Table 1.

The above constants and Equation 7 apply to the irrigation region below the "loading point".

When pressure drop curves for irrigated packing are examined, initially the liquid hold-up is solely dependent on the irrigation rate and the packing characteristics and this is shown by the irrigated pressure drop curves which parallel the dry bed pressure drop curve on the logarithmetric plot. As the gas velocity is increased beyond a critical value, the "loading point", the pressure drop rises more rapidly as the flow of liquid through the packing is impeded by the upward high velocity gas flow.

It is possible, from published data, to develop a correction method for the increase in pressure drop due to liquid loading and, based on the pressure drop calculated on the basis of no loading, to make a reasonable estimate of the pressure drop in the lower part of the loading region. When the calculated pressure drop rises well above one inch per foot, however, the tower may or may not flood, either locally or totally, such that a further extension of the proposed loading correction needs to be treated with caution. These comments apply to the large saddle packing, which is almost impossible to flood under industrial conditions. Massive acid carry-over to the mist eliminators would normally make a tower inoperable well before it floods. As is the case for the effect of irrigation, an exponential function is likely to best fit the data and a relationship can be formulated as follows.

Eq. (8)

and

Eq. (9)

The relationship between the square of the irrigated pressure drop and the correction factor for the loading region is based on data from the standard saddle but should apply in general. For the HP^TM and other saddles, the best estimate is that C₄ is equal to 1.6. For this numerical value of C₄, the pressure drop in the loading regime rises sharply as the calculated irrigated pressure drop rises beyond 0.6" W.C. per foot.

The complete CECEBE correlation for the irrigated packing pressure drop in large towers at a pressure drop of 1" W.C. per foot of packing or less can now be written as:

Eq. (10)

For clean HP^TM packing in a large tower, is 0.75.

Figure 5 is a plot of Equation 10 of pressure drop in a tower as a function of superficial gas velocity for HP^TM saddle packing and standard 3" saddle packing. For a given superficial gas velocity, say 5 ft/s, the savings is pressure drop in replacing standard 3" saddles with HP^TM saddles will exceed 50%. Alternatively, for a given pressure drop, say 0.4" W.C./ft, the superficial velocity can be increased by over 40% flow rate.

Figure 5: Plot of Equation 10 for Several Packings

This CECEBE HP^TM design correlation has been verified through measurements in a number of large acid towers operating with gas velocities in excess of 10 ft/s. Separate measurements were made for the pressure drop in the tower nozzles and packing support. It should be noted that data with pressure drop above 0.6" W.C. is very limited. At present, it is not recommended to design a new tower to operate in the loading zone. The proposed correlations are useful, however, to assess the effect of fouling and to analyze the performance of a tower which is asked to accommodate an increase in production.

Additional plant operating data would be very useful to confirm or adjust the constants used for the proposed correlation. Since the Florida fertilizer industry has many acid towers in operation, there is probably no better source for both, realistic evaluations of the packing pieces needed per unit volume, and the pressure losses under plant conditions. Where new packings are used, the background equations also offer an opportunity for comparison with other packings in similar sized towers.

The rise in pressure drop due to irrigation and loading can also be used to assess the associated liquid hold-up. As the gas velocity increases, the kinetic energy in the gas allows liquid to be entrained by the gas and to be carried up to the next packing layer, thereby increasing the "apparent" internal irrigation rate between packing layers. This mechanism will increase significantly the liquid hold-up and the gas velocity in the bed, and the pressure drop will rise as illustrated in the previous section. Operation in the loading regime is normally avoided because good design techniques for this region are lacking. With good pressure drop data, it is possible to evaluate this phenomenon is more detail and perhaps extend the design and operating limits of packed towers.

Since the increase in gas pressure drop is associated with a reduction in the void fraction by liquid hold-up, it is also possible, by using the basic equations, to calculate the actual void fractions for operating towers and deduce both hold-up of liquid and fouling with packing chips or sulfate. In this case, void fraction is the gas space free of packing, liquid hold-up, and fouling deposits. Having available pressure drop data, gas and liquid flow rates and gas density allows the calculation of from Equation 10.

Comparison of Packing

When new towers are designed,an allowance for fouling should be made by adjusting the void fraction using Equation 2, for example by assuming that the void fraction is lowered by fouling. Similarly, when repacking a tower, the old packing probably will have a lower void fraction due to fouling or shrinkage, than when it was originally installed.

Figure 6: Generalized Pressure Drop Correlation (Click to Enlarge)

The pressure drop and power saving associated with repacking as calculated by normal techniques will probably underestimate the benefits. Similarly, newer packings, for which full size tower data are available, can be more rigorously screened for economic viability. In many cases, the power savings from repacking with a lower pressure drop packing will often pay for packing premium costs in one or two years. If there are fewer chips, further savings will result, but these are harder to quantify. In general, the cheapest packing available is probably the worst bargain for the owner unless the towers had been grossly oversized to start with. This point has not yet been recognized by many owners.

Two-phase flow has been studied for many years, mostly in small towers. The common design methods advocated can generate significantly different solutions for the same packed tower duty depending on the methods used. Several different approaches to packed tower design were reviewed in preparing this paper.

Let us consider now the different methods of tower sizing in more detail. The most common approach is that described in Perry''s Chemical Engineers'' Handbook, 6th Edition. The GPDC is shown in Figure 18-38 in this reference and has been copied here from another reference as Figure 6. The packing factor "F" for Figure 6 are provided in Table 2. The data on which this figure is based were obtained in the fifties and sixties using a thirty inch diameter column. Packing factors are shown in a separate table. Sherwood, Holloway, and Eckert were the the primary contributors to this work but Eckert is responsible for the form as it appears in Perry''s. Initially the key characteristic of the packing was the ratio of the interfacial area (a) of the packing to the void fraction in the bed (a/³). Leva has also been given credit for this correlation, but he does not claim credit nor does he recommend it.{parse block="google_articles"}

It is now known that using data derived from large packing in small towers introduces errors both in the interfacial area, more saddles in a cubic foot than predicted, and lower void fraction. For the same reason, drastically revised packing factors for the large packings might bring the correlations in line but to our knowledge, this correction has not been proposed. Such correction will likely result in a doubling of the packing factor for three inch saddles.

Table 2: Packing Factors (F) -- Random Dumped Packings

A second method for design is to use gas pressure drop curves measured in water-air systems for the various packings and commonly included in packing brochures. A sample is shown in Figure 3. These graphs are specific to the packings in question. U.S. Stoneware provided a book of such curves, but unfortunately the three inch saddle had not been developed at that time so the only graphs in the Norton brochures are quite small. The data come again from small test towers. Adjustments of the data to larger tower is required.

A third proprietary approach follows more closely that suggested of Leva in calculating dry bed pressure drop with a correction factor for the effect of irrigation. The empty tower velocity head lost per unit height is calculated from the liquid flow per unit area and two proprietary constants. G. Morris of ICI, the designer of this technique, did not recommend using it with high gas flows and pressure drops and several

constants are involved for each packing size. In our review, the approach of Leva is simpler with the dry bed pressure drop calculation involving the empty tower liquid velocity and one constant for each packing. An exponential term including the liquid mass velocity in the empty tower is then used to correct for the impedance of gas flow by the liquid. The correction for acid flow will range as high as 100% in main absorbers with high liquid flows and 50% in drying towers and final absorbers. Morris adjusted his void fractions to towers with diameters twenty to thirty times the packing size which corresponds to six to eight feet, a step in the right direction, though not far enough.

Many engineers, over the years, have developed correlations and programs for packing pressure drop from the published test tower data and used the correlations to predict pressure drop. One such program, generally made available by a supplier of packings, has been used to evaluate pressure drop for the test case for both standard 3" saddle packing and a structured ceramic packing.

Several comments need to be made about the data used in these evaluations. The first is that the number of pieces of packing in a unit volume of tower vary significantly with the ratio of tower diameter to packing size as was shown in Figure 2. Essentially all of the data used in the correlations were developed in towers in the 2.5 to 3 foot range. When the packing was small such as 1", the number of pieces per

unit volume was reasonably high and usable in large towers. With large saddles, however, the packing density was much lower than found in full size towers and the pressure losses were much lower. In packing supply quotes, the "so-called" settling allowance is simply a correction factor to ensure there are enough pieces of packing to fill the tower. In many cases know to the authors, packing purchased to refill a tower of specified volume was found to be insufficient to physically fill the tower, even though the exact tower dimensions were given to the supplier. A "shrinkage" of 1.5 ft out of a 12 ft packing height was not unusual, and often caused a very embarrassing start-up situation. It is on the basis of this experience that CECEBE went through the trouble of making a saddle count when packing a number of new towers. Similar data on standard saddles would also be very useful. It is also suggested that the quantity of packing normally supplied by manufacturers to fill a cubic foot of tower will actually normally not fill a cubic foot in a large tower and owners might have a legitimate case for a better definition of what is needed to fill a cubic foot of tower volume.

The tower diameter for the standard three inch saddle using the different design techniques are shown in Table 3 while the diameters for different packings are shown in Table 4.

What the above tables show is that various methods available for tower sizing give widely different answers (Table 3) and that there is significant capacity difference between different packing (Table 4). This capacity difference can be exploited when a plant is being expanded by repacking with higher performance packing.The Economic Cost of Pressure Drop

Packing performance can be assessed in terms of mass transfer but can also be assessed in terms of the power needed to push the gas through the tower. A more open tower with lower pressure drop will offer greater production capacity or less power consumption or less acid mist or a combination of all three. For the 2000 TPD case chosen, the overall plant pressure drop typically will be around 220" W.C. {parse block="google_articles"}(atmospheric pressure is 406.8" W.C.). The pressure saving from high performance packing in one tower would amount to around 70 kW for a pressure drop saving of 5" W.C. This is worth between fifty and sixty thousand dollars per year based on $0.10/kW hr. For standard packing, around 3300 ft³ would be needed to fill the tower which, with an allowance for settling, would cost around sixty thousand dollars. The HP^TM saddle will sell for a premium cost due to its more expensive manufacturing process and structured packing will cost still more. With the HP^TM saddle, the power saving would pay for the extra packing cost in less than a year. A second important issue is the fact that, by using HP^TM packing, in a tower repack case, additional acid production will become possible, not only because of its inherent lower pressure drop, but also because of its low breakage during installation. In one 13 ft acid tower with a bottom chimney screen, less than half a bucket of packing chips were collected during the first annual turn-around.

Discussion of Results

Table 3 shows that there is a wide discrepancy in tower diameter prediction when different sizing methods are used. From practical experience and discussion with colleagues, the GPDC approach appears sound but the packing factors published for large packings appear questionable. This is likely due to the use of the "void fraction effect" in small pilot towers. An adjustment of packing factors for all three of the larger size saddles would appear to be in order and could result in at least a doubling of the packing factors for the 3" saddle. Possibly the published packing densities of Figure 2 and the basis Leva equation would give sufficient guidance. The two proprietary programs gave similar results which suggested a similar logic basis but the details of the program were not available. Again, these programs are only as good as are the data used in the correlations. The question is, are these data obtained from a pilot tower or a full scale tower.{parse block="google_articles"}

The last design approach listed in Table 3 is based on data collected by CECEBE and NORAM in a full size tower using the actual void fractions. There is some concern about designing to relatively high pressure drops as with any correlations when the pressure drop approaches 0.75" W. C. There is no benefit to the owner if a supplier is overly optimistic on packing performance only to have the owner ultimately suffer the consequences. Nevertheless, higher quality packing will give measurable and economically quantifiable performance advantages.

In Table 4, the void fractions have been adjusted to larger tower diameters and a number of packings have been evaluated on the basis of the CECEBE design techniques. As can be seen, larger standard saddles give smaller tower diameters just as one would expect. For the HP^TM saddle, field data are available and a structured ceramic packing, for which a proprietary program was available, is also listed. Interestingly, the diameter associated with the HP^TM saddle was essentially the same as that predicted from several of the programs for the standard saddle and well below that which one can expect for the standard 3" saddle. The structured ceramic program predicted a slightly smaller diameter but it is not backed by any published field data. However, the nature of this structured packing suggests that it is not significantly affected by tower size. As mentioned already, its major disadvantage is that it requires special attention to liquid distribution. No capacity benefit is gained if liquid distribution is implemented through a layer of standard 3" saddles.

If one is considering installation of new towers, the cost of the tower is one issue. The cost of packing is a second. Standard packing is sold normally on the basis of a definition of a piece density which falls short of what is needed to cover a "settling allowance". Often fifteen percent extra packing may be needed. Combining this with a large diameter tower can result in as much as fifty percent more packing being needed over more recent high performance packing to get an equivalent result. Combining this with the cost of the tower shell, sound economics would suggest that the best packing will probably give the best economic solution and the cheapest packing, possibly the worst solution. The difficulty has been the wide discrepancy in the ability to predict pressure drop in large packed towers. In several cases recently, HP^TM saddles have been supplied to owners who were concerned about pressure drop and power consumption and could justify the expense on the power saved. With rising power costs, this incentive will become even more powerful. The discussion above on pressure drop offers data relevant to this issue.

Further Work, References,and Acknowledgements

Further Work

Additional work on towers with standard and new high performance packings is called for. In our view, the approach of Leva has many merits. Data from fertiziler plants would be very useful. Characterization of newer packings would also allow better comparisons which would help owners in deciding what to do. There is also an open question as to packing heights and the extent to which mist eliminators can be counted on to finish the mass transfer process. Similarly, using a large number of irrigation points has been advocated by some to allow much shorter packing heights. Candles can do a good job of removing residual sulfur trioxide but only if the candles are irrigated. This happens sometimes by default if the distributor generates spray if the tower is overloaded. Candles are more expensive than mesh pads and it may be less expensive to use more packing. More irrigation points, over one point per square foot, would clearly be desirable with smaller packing and also with mesh or similar packings where liquid tends to fall {parse block="google_articles"}vertically. The value of high irrigation points counts with large saddles is debatable, especially as more points leads to systems which are more expensive and also much more easily plugged with chips because of smaller distributors holes or downcomers.

References

1. VFF Data Sheet for Novalox Saddles

2. U.S. Stoneware Design Manual, Pressure Drop vs. Gas Rate for 1 1/2" INTALOX Saddles, Ceramic

3. Dr. M. Leva, private communication

4. Tower Sizing Program, Kock Engineering

5. Dr. G. A. Morris, ICI, private communication

6. Perry''s Chemical Engineers'' Handbook, 6th Edition, Figure 18-38

7. Website, http://www.epa.gov/, TTNWeb, Clearinghouse for Inventories and Emmission Factors

Acknowledgements

Wave Packing is a registered trademark of MEC

INTALOX is a registered trademark of Norton Chemical Company

NOVALOX is a registered trademark of VFF Industries

HP^TM is a registered trademark of CECEBE Technologies Inc.

Introduction{parse block="google_articles"}

One way to cut the cost of chemicals is to utilize microorganisms to break down most of the hydrogen sulfide using oxygen from the air.

Using microorganisms to remove odor or volatile organic compounds from air streams is not a new idea. Biofiltration has been used, especially outside the United States, for many years.^1,2 For H₂S odor control, the key is to provide an ideal habitat for the growth of sulfide-oxidizing bacteria, to the exclusion of competing microbes which normally predominate in aerobic treatment processes.

Several species of microorganisms can oxidize hydrogen sulfide to form odorless sulfuric acid. A few species of the genus Thiobacillus are capable of oxidizing H₂S at low pH. Thiobacillus thiooxidans, in particular, thrives at pH <3, and its growth is not inhibited until the pH falls below 1.³

Figure 1: Conventional Biofilters6

Efficient removal of H₂S requires media with enough surface area to maintain a large population of sulfide-oxidizing microbes.

Porous media such as soil, peat, compost and/or wood chips work well in biofilters for removal of organic vapors, although they require careful control of temperature and humidity. If the air is not fully saturated with water vapor, some of the medium may dry out, inactivating the microbes on it. On the other hand, excessive moisture can cause water to accumulate in the media, and eventually wash away nutrients.

The weight of moist media limits the depth that can be used without excessive compression. Worst of all, sulfuric acid formed by biological oxidation of sulfur compounds can degrade such media, causing them to collapse.

As a result, conventional biofilters using these media often need a caustic scrubber as a pretreatment stage to humidify the inlet air stream and remove sulfur compounds.

To eliminate the need for pretreatment of the air or for periodic replacement of degraded media, biofilters can be built using acid-resistant inorganic substrates such as porous lava rock. These are referred to as "trickling biofilters" (or "biotrickling filters") because the media is kept wet-regardless of the humidity-by continuous circulation of water.

Figure 2: Trickling Biofilter

However, the weight of rock media makes it difficult to handle, and limits the depth of a filter bed that can be installed without expensive reinforced structures. The fan power needed to force air through a bed of lava rock is also quite high. As a result, trickling biofilters using rock media must be sized for very low gas velocities, resulting in huge footprints.

Nitrifying trickling filters using plastic media have also been used for H₂S odor control,⁴ but the limited surface area of conventional trickling-filter media results in relatively low bacterial populations per unit volume. Air residence times on the order of minutes are required for efficient odor removal. The cost of such enormous filters cannot be justified unless they are needed to nitrify wastewater.

In order to overcome the drawbacks of conventional media, Lantec Products has developed a high-density polypropylene media known as HD Q-PAC^.

This media is acid-resistant, lightweight, easy to handle, and rigid enough to walk on. It can be stacked to any desired depth. It provides 132 ft² of plastic surface per cubic foot, yet it has a high void fraction, so that even when coated with a layer of biofilm it still presents much less resistance to air flow than compost or rock media.

This makes it possible to treat air at higher superficial velocities with reasonable fan power requirements, so trickling biofilters can be made taller rather than wider, saving valuable space in crowded treatment facilities.

Table 1: Physical Properties

	Material:	Polypropylene
	Specific Surface Area:	132 ft2 / ft3
	Drip Points:	75,000 / ft3
	Bulk Density;	7.5 lb / ft3
	Void Fraction:	87.8 %
	Smallest Grid Opening:	0.16" 0.16"
	Standard Module Size:	12" x 12" x 12"
Figure 3: HD Q-PAC®

Test Installations

Wastewater Pump Station Odor Control

{parse block="google_articles"}HD Q-PAC^® was used in a small trickling biofilter as the only odor removal system at a pump house in Saco, Maine. The pump house is situated in the middle of a residential area, with the closest residence is no more than 20 feet away. The treatment plant had gotten odor complaints from residents every summer. A biotrickling filter was installed in May 1999 to evaluate its effectiveness in reducing high H₂S levels during the warm summer months.

Two 55-gallon liquid storage drums with inside diameter of 20 inches were welded together and used as the biofilter vessel. HD Q-PAC^® was installed in the tower with the needles oriented vertically. Gaps between the media and the walls of the drums were filled with separated pieces of HD Q-PAC^®.

Figure 4: Test Unit at Saco, Maine

A centrifugal blower rated at 80-100 cfm pushes the air into the biofilter. Contaminated air enters the bottom of the vessel, passes upward through the biofilm-coated media, and exits through the top. H₂S inlet concentrations in the pump house air range from 1 to 90 ppm_v.

Water is recirculated at 6 gpm. Fresh water with nutrients is added at 1 gpm for 15 min every day. H₂S concentrations are measured using a Scott Alert Meter (Model S108) which is calibrated monthly.

When the biofilter was started up in May 1999, the H₂S removal efficiency increased gradually as biofilm developed on the media, but it remained below 45%. A commercial lawn-fertilizer dispenser was connected to the make-up water line to add ammonium phosphate and urea as micronutrients. After that, the H₂S removal efficiency rose to over 90% within a few days. Since then, the removal has been consistently over 90%, even though the inlet H₂S level varied as much as 400% within the same day.

Throughout the summer of 1999, with recording-breaking high temperatures, the treatment plant did not receive a single complaint about pump station odors.

Figure 5: Hydrogen Sulfide Removal at Windy Hill Pump Station in Saco, Maine

Central Wastewater Treatment Plant

HD Q-PAC was also tested in a trickling biofilter to remove hydrogen sulfide from exhaust air at the Hyperion Treatment Plant in Los Angeles, California.

This test filter has an inside diameter of 4.5 ft and is packed with 7 ft of media. HD Q-PAC^® was installed in the tower with its needles oriented horizontally. Each rectangular module of media was stacked tightly against the others, leaving no gaps between them. Gaps between the HD Q-PAC and the walls of the circular tower were filled in with small pieces of porous rock.

Figure 6: Hyperion Plant Test Unit

A blower sends untreated air into the bottom of the tricking biofilter. The air flows upward through the biofilm-coated media, while the water trickles down over it. The treated air exits the top of the unit.

The filter was initially used to treat 700 cfm of air containing 2-20 ppm_v of H₂S. Water was recirculated over the media at a rate of 10 gpm.

The unit was started up by filling the 300-gal sump with secondary effluent from the treatment plant, then running the fan and recirculation pump continuously until bacteria began to colonize the media, and the pH of the water decreased to less than 2.0. After that, a portion of the acidic solution was made to overflow every 4 hours by adding secondary effluent at 3 gpm for 20 minutes.

In addition to controlling the pH, the 360 gallons of make-up water added each day provided micronutrients needed for growth of the biofilm. (Thiobacillus thiooxidans is autotrophic; it uses atmospheric CO₂ as its carbon source.)

The H₂S concentrations in the inlet and outlet air streams were measured daily using an Interscan Voltammetric Sensor.

The removal efficiency of H₂S increased steadily for the first few days of operation, reaching 90% within 10 days.

Figure 7: The Author Installing HD Q-PAC®

Ever since then, the removal efficiency has remained between 90% and 95%, with higher efficiencies recorded occasionally. Since the initial start-up period, the H₂S removal efficiency has never fallen below 90%.⁵ This is in a small trickling biofilter with less than 10 seconds of residence time.

Parametric studies aimed at optimizing the operating conditions are now under way, and will be reported in a future paper. However, the consistent performance of this test unit over a period of months has demonstrated conclusively that the proper environment for Thiobacillus growth can be maintained using HD Q-PAC^® and extremely simple equipment.

Possible Applications

The biofilter at Hyperion Treatment Plant using HD Q-PAC^® will be scaled up to pretreat large volumes of exhaust air which is now being processed by conventional wet scrubbers.

Wastewater treatment plants in urban areas are among the world''s largest consumers of sodium hypochlorite. {parse block="google_articles"}By removing 90% or more of the H₂S using atmospheric oxygen, the operating cost of chemical oxidants for the scrubbers can be cut by hundreds of thousands of dollars per year. The existing scrubbers will continue to function as a "polishing" stage, and as a back-up in case of any problems with the trickling biofilters.

These biofilters are particularly well suited for odor control at isolated pumping stations and other facilities where there is no-one to operate a conventional wet scrubber, even if a water system could afford the equipment and the chemicals needed to scrub small air streams at many scattered locations. These filters are simple enough to run automatically without operator attention, and with no need to store hazardous chemicals at multiple unguarded sites.

In many developing countries, the capital and operating costs of wet scrubbers are more than treatment plant budgets can bear. The simplicity of trickling biofilters, and their ability to operate without expensive chemicals, provide a badly needed alternative in this situation.

Trickling biofilters may also find use as simple pretreatment stages for conventional biofilters for VOC removal. They can humidify air and greatly reduce its sulfur content, extending the useful life of water-absorbent biofilter media while eliminating the need for treatment chemicals.

References

1.Bohn, H., "Soil and compost filters for malodorous gases," J. Air Pollution Control Assoc. 25, p.953 (1975).

2. Ottengraf, S. and Van Den Oever, A., "Kinetics of organic compound removal from waste gases with a biological filter," Biotechnol. Bioeng., 25, p. 3089, (1983)

3. Devinny, J., Deshuesses, M., Webster, T., "Biofiltration for Air Pollution Control," Lewis Publishers, Boca Raton, p. 74, (1999).

4. Lutz, M. and Farmer, G., "Pulling double duty: A Colorado plant''s trickling filters treat odor while reducing wastewater nitrogen content," Water Environment Federation Operations Forum, 16 (7), pp.10-17 (1999)

5. Steve Johnson, Hyperion Wastewater Treatment Plant, Los Angeles, California (personal communication)

6. Devinny, J., Deshuesses, M., Webster, T., "Biofiltration for Air Pollution Control," Lewis Publishers, Boca Raton, p. 9, (1999).

Individuals with different backgrounds, from engineering to biochemistry, can make significant contributions to the understanding of biosorption. Interdisciplinary efforts are essential to exploit this technology commercially. A chemical engineering background is particularly useful for expanding the application of this technology in large-scale process industries.

Threats from the Environment

The greatest demand for metal sequestration today comes from the need to immobilize the metals released to the environment (or mobilized) by and partially lost through human technological activities. It has been established that dissolved metals (particularly heavy metals) escaping into the environment pose a serious health hazard¹¹. They accumulate in living tissues throughout the food chain (Figure 1), which has humans at its top, multiplying the danger. Thus, it is necessary to control emissions of heavy metals into the environment.

Am example of one method for prioritizing the recovery of ten metals is presented in Table 1. This may be simplistic, but it provides a useful direction by ranking metals into three general priority categories:

1. Environmental Risk (ER)
2. Reserve Depletion Rate (RDR)
3. Combination of ER and RDR.

Environmental risk assessment could be based on a number of different factors, which could also be weighted.

The Need for Novel Technology

Conventional techniques to remove toxic metals and radionuclides, such as ion exchange and precipitation, lack specificity and are ineffective at low metal ion concentrations. The need for effective and economically viable technologies is driven by environmental pressures such a:

• Stricter regulations with regard to the metal discharges are being enforced, particularly in industrialized countries.
• Toxicology studies confirm the dangerous impacts of heavy metals.
• Current technologies for the removal of heavy metals from industrial effluents often create secondary problems with metal-bearing sludge.

Biosorption Mechanisms

Various metal-binding mechanisms have been postulated to be active in biosorption, such as:

• Chemisorption by ion exchange, complexation, coordination and/or chelation{parse block="google_articles"}
• Physical Adsorption
• Microprecipitation
• Oxidationeduction.

Due to the complexity of the biomaterials used, it is possible that at least some of these mechanisms are acting simultaneously to varying degrees, depending on the biosorbent and the solution environment.

Ion Exchange

Ion exchange is a reversible chemical reaction wherein an ion in a solution is exchanged for a similarly charged ion attached to an immobile solid particle. These solid ion-exchange particles are either naturally occurring inorganic zeolites or synthetically produced organic resins. Synthetic organic resins are the predominant type used today because their characteristics can be tailored to specific applications.

Ion exchange reactions are stoichiometric and reversible, and as such they are similar to other solution-phase reactions. For example, in the reaction

NiSO₄ + Ca(OH)₂ Ni(OH)₂ + CaSO₄

the nickel ions of the nickel sulfate (NiSO₄) are exchanged for the calcium ions of the calcium hydroxide Ca(OH)₂ molecule.

Chelation

The word chelation is derived from the Greek word chele, which means claw, and is defined as the firm binding of a metal ion with an organic molecule (ligand) to form a ring structure. The resulting ring structure protects the mineral from entering into unwanted chemical reactions. Examples include the carbonate (CO₃^2-) and oxalate (C₂O₄^2-) ions:

Figure 1: Carbonate and Oxalate Complexes

Coordination (Complex Formation)

A coordination complex is any combination of cations with molecules or anions containing free pairs of electrons. Bonding may be electrostatic, covalent or a combination of both; the metal ion is coordinately bonded to organic molecules. Example of the formation of a coordination compound are:

Cu²⁺ + 4H₂O [Cu(H₂O)]₄²⁺

Cu²⁺ + 4Cl^- [CuCl₄]^2-

where coordinate covalent bonds are formed by donation of a pair of electrons from H₂O and Cl^- (Lewis bases) to Cu²⁺ (Lewis acid).

In general, biosorption of toxic metals and radionuclides is based on non-enzymatic processes such as adsorption. Adsorption is due to the non-specific binding of ionic species to polysaccharides and proteins on the cell surface (Figure 2) or outside the cell^4,19. Bacterial cell walls and envelopes, and the walls of fungi, yeasts and algae, are efficient metal biosorbents that bind charged groups. The cell walls of gram-positive bacteria bind larger quantities of toxic metals and radionuclides than the envelopes of gram-negative bacteria.

Bacterial sorption of some metals can be described by the linearized Freundlich adsorption equation:

log S = log K + n log C

where:

S is the amount of metal absorbed in mol/g

C is the equilibrium solution concentration in mol/L

K and n are the Freundlich constants.

Biomass deriving from several industrial fermentations may provide an economical source of biosorptive materials. Many species have cell walls with high concentrations of chitin, a polymer of N-acetyl-glucosamine that is an effective biosorbent.

Biosorption uses biomass raw materials that are either abundant (e.g., seaweeds) or wastes from other industrial operations (e.g., fermentation wastes)⁹. The metal-sorbing performance of certain types of biomass can be more or less selective for heavy metals, depending on the type of biomass, the mixture in the solution, the type of biomass preparation, and the chemical-physical environment.

It is important to note that the concentration of a specific metal in solution can be reduced either during the sorption uptake by manipulating the properties of the biosorbent or upon desorption during the regeneration cycle of the biosorbent.

Sources of Biomass for Biosorption

Sources of biomass include:

• Seaweeds
• Microorganisms (bacteria, fungi, yeast, molds)
• Activated sludge
• Fermentation waste{parse block="google_articles"}
• Other specially propagated biomasses.

Biosorbents must be hard enough to withstand the application pressures, porous and/or "transparent" to metal ion sorbate species, and have high and fast sorption uptake even after repeated regeneration cycles¹⁷.

Granulation of biomass materials into suitable cost-effective biosorbents is a crucial step for the successful application of biosorption processes.

The objectives of granulation are to:

• Establish the behavior of native biomass in a packed-bed reactor
• Establish the effectiveness of biomass granulation and reinforcement
• Determine the effect of size reduction on sorption capacity
• Determine the feasibility of biomass processing.

Conventional granulation technologies are rather advanced, and their adaptation will likely yield desirable biosorbent granules⁵. Because of the wide variety of biomass types, extensive experimentation will undoubtedly be required.

The need to transport raw biomass may also present some logistical problems. Microbial biomass has a high water content and is prone to decay, so drying may be required if it cannot be processed and/or granulated directly on location in the wet state.

Equilibrium Modeling

Biosorption has been studied as simplified sorption systems, usually containing one heavy metal. This is an appropriate simplification for effective experimentation.

Table 2 summarizes some of the simple sorption isotherm models that are most frequently applied. A particular model may not apply to a particular situation, and in some cases more than one model may explain the biosorption mechanism. There is no critical reason to use a more-complex model if a two-parameter model (such as the Langmuir and Freundlich isotherm models) can fit the data reasonably well.

Table 2: Frequently used single-component adsorption models¹¹

Isotherm	Equation	Advantages	Disadvantages
Langmuir		Interpretableparameters	Not structured; Monolayer sorption
Freundlich		Simpleexpression	Not structured
Combination of Langmuir and Freundlich		Combination of the above two	Unnecessarily complicated
Radke and Prausnitz		Simple expression	Empiral; Requires three parameters
Redlich Peterson		Approaches Freundlich at higher concentrations	No significant advantages
Braunauer, Emmer, and Teller (BET)		Multilayer adsorption	------
Dubinin-Radushkevich		Temperature dependent	Behavior is not limited in the Henry''s Law regime

As a matter of practicality, multi-metal biosorption models such as those in Table 3 must be used judiciously.

Table 3: Frequently Used Multi-Component Adsorption Models¹¹

Isotherm	Equation	Advantages	Disadvantages
Langmuir (Multi-component)		Constants have physical meaning; isotherms level off at maximum saturation.	Not structured; doesn''t reflect the mechanism well
Combination (Langmuir and Freundlich)		Combination of Langmuir and Freundlich	Unnecessarily complicated

where:

C_e is the equilibrium solute concentration in the fluid

K,n are the Freundlich isotherm constants

a_i,b_i are the Langmuir isotherm parameters

e is the column bed porosity; Polanyi''s adsorption potential

q_m is the Langmuir maximum metal uptake in mg/g

W_o,W_m are the initial and final volumes, respectively, in L

K_{(with various subscripts)} are the intrinsic equilibrium constants

Q is the meta uptake in mg/g

b is the Polanyi scaling factor in the Polanyi models.

The sorption uptake, q, can be expressed in different units depending on the purpose of the exercise:

• For practical and engineering process evaluation purposes eventually concerned with process mass balances, it is customary to use weight per (dry) weight (e.g., mg of metal sorbed per gram of the (dry) sorbent material).
• Ultimately, mainly because of reactor volume considerations (e.g., a packed-bed column), the uptake may also be expressed on a per volume basis (e.g., mg/L). However, the porosity may complicate the quantitative comparison of biosorption performance.
• Only when working on the stoichiometry of the process and when studying the functional groups and metal-binding mechanisms might it be useful to express q on a molar or charge equivalent basis - again, per unit weight or volume of the sorbent (e.g., mmol/g or mequiv/g).

It is relatively easy to convert among these units; the only problem may arise with the sorbent weight-volume conversions. For scientific interpretations, the sorbent material dry-weight basis is thus preferred.

The use of "wet biomass weight" should be discouraged, unless the wet-weight-to-dry-weight conversion is well specified. Different biomass types are likely to retain different moisture contents, intracellular as well as that trapped in the interstitial space between the cells or tissue particles (e.g., seaweed particles). Different types of biomass obviously compact in a different ways. When centrifuging biomass, the g-force and time need to be specified, and even then it is difficult to make any comparisons. All this makes the "wet biomass weight" citation very approximate at best and generally undesirable.

Biomass Types

The assessment of the metal-binding capacity of some types of biomass has gained momentum since 1985¹⁸. Indeed, some biomass types are very effective in accumulating heavy metals.

Availability is a major factor to be taken into account to select biomass for clean-up purposes. The economics of environmental remediation dictate that the biomass must come from nature, or even be a waste material. Seaweeds, molds, yeasts, bacteria, and crab shells, among other kinds of biomass (Figure 3), have been tested for metal biosorption with very encouraging results.{parse block="google_articles"}

Biosorpents

Some biosorbents can bind and collect a wide range of heavy metals with no specific priority, whereas others are specific for certain types of metals. When choosing the biomass for metal biosorption experiments, its origin is a major factor to be considered.

Biomass can come from:

• industrial wastes which should be obtained free of charge
• organisms that can be obtained easily in large amounts in nature (e.g., bacteria, yeast, algae)
• fast-growing organisms that are specifically cultivated or propagated for biosorption purposes (crab shells, seaweeds).

Organisms for Biosorption

There is a wide variety of microorganisms (Table 4), including bacteria, fungi, yeast, and algae, that can interact with metals and radionuclides and transform them through several mechanisms.

Cost-effectiveness is the main attraction of metal biosorption. This cost-effectiveness can be maintained by using the microbial biomass directly where possible. In addition, biosorbents derived from microbial biomass through a simple process are expected to be the lowest-priced and most-economical for metal removal.

It has been suggested that numerous chemical groups contribute to biosorption metal binding, by either whole organisms such as algae and bacteria or by molecules such as biopolymers. These include hydroxyl, carbonyl, carboxyl, sulfhydryl, thioether, sulfonate, amine, imine, amide, imidazole, phosphonate, and phosphodiester groups. The importance of any given group for biosorption of a certain metal by a certain biomass depends on such factors as the number of sites in the biosorbent material, the accessibility of the sites, the chemical state of the sites (i.e., availability), and the affinity between the site and the metal (i.e., binding strength). For covalent metal binding, even an occupied site is theoretically available; the extent to which the site can be used by a given metal depends on its binding strength and concentration compared to the metal already occupying the site.

Some types of industrial fermentation waste biomass are excellent metal sorbers. It is necessary to realize that some "waste" biomass is actually a commodity, not a waste. This applies particularly to the ubiquitous brewer''s yeasts sold on the open market, usually as animal fodder.Activated sludge from wastewater treatment plants has not demonstrated high enough metal-sorbing capacities. Some types of seaweed biomass offer excellent metal-sorbing properties, and sometimes a local economy can benefit from turning seaweeds into a resource.

As a fallback, biomass with a high metal-sorbing capacity can be specifically grown relatively cheaply in fermenters using low-cost or even waste carbohydrate-containing growth media such as molasses or cheese whey.

Experimental Sorption Isotherms

It is relatively easy to obtain equilibrium sorption data for a single sorbate in the laboratory. A small amount of the sorbent is brought into contact with a solution containing the sorbate of interest. The conditions of the sorption system, particularly pH, must be carefully controlled at the required values over the entire period of contact until the sorption equilibrium is reached. This may take a few hours or much longer, depending on the size of the sorbent particles and the time it takes until they attain sorption equilibrium.{parse block="google_articles"}

A simple preliminary sorption kinetics test will establish the exposure time necessary for the given sorbent particles to reach the equilibrium state. The following procedure provides an example for obtaining the experimental sorption equilibrium data points for the isotherm:

1. Prepare the sorbate in solution at the highest concentration of interest.
2. Prepare dilutions covering the entire concentration range, from 0 (blank) to the maximum.
3. Adjust the conditions, e.g., pH, ionic strength, etc.
4. Determine the sorbate initial concentrations (C_i ) in all the liquid samples.
5. Distribute the samples into containers of appropriate volumes (30-150 mL of liquid) such as flasks or test tubes; prepare samples in duplicate, triplicate or as required.
6. Accurately weigh each quantity of the biosorbent solids to be used in the tests and record the weights (S, mg). It may help to be able to roughly estimate the anticipated sorption uptake so that there is an easily detectable final sorbate concentration in each sample solution at equilibrium. If too much sorbent is added, there may be virtually no sorbate left in the solution, precluding a reliable analysis. Varying the initial concentration could cause the sorbent weight to fluctuate, which has to be precisely known for each sample. Metal depletion in the solution must be avoided because it renders such samples useless.
7. Add the sorbent solids into each sample solution and provide rather gentle mixing over the contact period.
8. Make sure the conditions (especially pH) are controlled at constant values during the contact period. Use an appropriate acid or base for this; do not dilute the sorption system by adding excessive volume.
9. At the end of the contact period, separate the solids from the liquid by decantation, filtration, centrifugation, etc.
10. Analyze the liquid portion to determine the residual final sorbate concentration (C_f).
11. Calculate the sorbate uptake: q = V [L] (C_i - C_f) [mg/L] / S [g]. Note that q could also be determined directly by analyzing the separated solids and thus closing the material balance on the sorbate in the system. However, this usually presents analytical difficulties (digestion-liquefaction of solids, and/or very sophisticated analytical methods may be required).
12. Plot the sorption isotherm q vs. (C_f).

Comparison of Sorption Performance

The performance of sorbing materials needs to be evaluated and often compared. The simplest situation is when there is only one sorbate species in the system, in which case it is best to base the single-sorbate sorption performance on a complete single-sorbate sorption isotherm curve.

To fairly compare two or more sorbents, the comparison must be done under uniform conditions. These may be restricted by the environmental factors under which sorption may have to take place (pH, temperature, ionic strength, etc.), which may not necessarily be easily or widely adjustable. In particular, it is important to compare sorption performance under the same pH conditions, since isotherms can vary with pH.

The performance of the sorbent is usually gauged by its uptake (q). Sorbents can be compared based on their respective maximum uptake values (q_max), which can be calculated by fitting the Langmuir isotherm model to the actual experimental data (if it fits). This approach is feasible if q_max reaches a plateau. Some isotherms might not exhibit the asymptotic plateau represented by the Langmuir equation.

In general, one is looking for a "good" sorbent with a high sorption uptake capacity (q_max). Surface area in biosorption is not particularly important.

Types of Biosorption

Biosortpion can be carried out as a batch process, a continuous process, or a two-stage process with continuous metal recovery.

Biomass should be defrosted and washed with deionized water. To ensure equal quality of the biomass during all experiments, different kinds of biomass should be mixed together to obtain a uniform mixture.

Batch Process

Batch biosorption experiments can be done in a stirred vessel (Figure 4) with a working volume of approximately 100 mL. {parse block="google_articles"}A minimal amount of concentrated solution of Pb(NO₃)₂ (metal) can be added into a suspension of fungal pellets in water of various concentrations (25, 50, 100, 150 and 200 g of wet biomass per L of biomass suspension) to produce the desired initial metal concentrations of 10, 20, 50, 100 and 300 mg/L Pb²⁺ (metal). The decreasing metal concentration can be recorded as a function of the initial metal concentration and the biomass loading.

Continuous Process

Continuous process experiments can be carried out in a glass column having an inner diameter of 5-8 cm and filled with a packed bed of biomass pellets of varying heights (20, 40 and 55 cm), set with an adjustable plug (Figure 5). The effluent solution of metal ions can be fed from the top of the column with the help of a pump using varying flowrates. An inert bed of glass spheres can be placed at the bottom of the column below the active biomass bed to ensure homogenous distribution of the feed. The remaining metal concentration can be measured online in the effluent at the top of the column. The breakthrough curves can be recorded as a function of the flowrate and bed height.

Measurements of metal ion concentrations in the solution can be made online with metal-detecting electrodes or ion-selective electrodes, and may be verified with an atomic absorption spectrometer.

Two-Stage Process with Continuous Metal Recovery

Two-stage continuous biosorption and metal recovery can be carried out as shown in Figure 6. This process is similar to continuous biosorption, although the metal solution is adsorbed in two stages. After initial adsorption and filtration in stage one, the effluent is fed with fresh biosorbent into stage two, where further biosorption of the metal ion takes place. The effluent from the second stage is filtered to recover the metal ions and biosorbents. The effluent sample can be analyzed using an ion meter or by adsorption spectroscopy.

Desorption

Regeneration of loaded biosorbent is critical to keeping costs down and to recovering the metal(s) extracted from the liquid phase. The deposited metals are washed out (desorbed) and the biosorbent is regenerated for another cycle. The desorption process should result in:

• high-concentration metal effluent
• undiminished metal uptake upon re-use
• no physico-chemical damage to the biosorbent.

The desorption and sorbent regeneration studies might require somewhat different methodologies, beginning with screening for the most effective regenerating solution.

Because different metal ions have different affinities for the biosorbent, the uptake has some degree of metal selectivity. The selectivity of the elution-desorption operation may be different, which may serve as another means of eventually separating metals from one another if desirable.

The concentration ratio (CR) is used to evaluate the overall concentration effectiveness of the whole sorption-desorption process:

Obviously, the higher the CR, the better the overall performance of the sorption process, making the eventual recovery of the metal more feasible with higher eluate concentrations.

Recovery of the metal from these concentrated desorption solutions is carried out in a different plant by electrowinning. Following desorption of the metal(s), the column may be further pre-treated (e.g., pre-saturated with protons such as Ca, K, etc.) for optimum operation in the next metal uptake cycle. The specific type of pre-treatment used to optimize the column performance may vary.

Feasibility of Biosorption

For successful application on a large scale, any operation needs to be economically viable. The feasibility of a biosorption process depends on such factors as:

• biosorbent uptake performance
• the source of the raw biomass
• biomass granulation and treatment
• the desorption and regeneration processes used.

Often, the source of the biosorbent has a major impact on the feasibility of the operation. Biosorbents (biomass) should always be obtained from the least-expensive source, such as from the effluent of a fermenter, seaweeds from nearby bodies of water, algae, etc. The spent biosorbents can be regenerated at very low cost using water, so the material can be reused many times. Hence, considering the overall unit operations involved in biosorption, we can conclude that the process is generally economically viable.

Advantages of Biosorption

Biosorption is highly competitive with the presently available technologies like ion exchange, electrodialysis, reverse osmosis, etc. Some of the key features of biosorption compared to conventional processes include:

• competitive performance
• heavy metal selectivity
• cost-effectiveness
• regenerative
• no sludge generation.

Biosorption is particularly economical and competitive for environmental applications in detoxifying effluents from, for example:

• metal plating and metal finishing operations
• mining and ore processing operations
• metal processing
• battery and accumulator manufacturing operations
• thermal power generation (coal-fired plants in particular)
• nuclear power generation.

Conclusions

There appear to be many modes of non-active metal uptake by microbial biomass. Any one or a combination of them can be functional in immobilizing metallic species on biosorbents. A number of anionic ligands participate: phosphoryl, carbonyl, sulfhydryl and hydroxyl groups can all be active to various degrees in binding the metal.{parse block="google_articles"}

Many scientific studies are currently underway to provide a deeper understanding of biosorption and to support its effective application. Some pollution seems inevitable, and one might wonder what should be done to minimize it. Human populations need methods and technologies to clean waters and diminish the environmental dangers related to technological progress. Biosorption can be one such solution to clean up heavy metal contamination.

References

1. White, S.K. J. Am. Water Works Assoc. 1983, 75, 374.

2. Laul, J.C. Radioanal. Nucl. Chem. Articles 1992, 156, 235.

3. Benedict, B.; Pigford, T.H.; Levi, H.W. Nuclear Chemical Engineering, McGraw-Hill: New York, NY 1981.

4. Volesky, B.; Tsezos, M. U.S. Patent 4320093, 1981. Canadian Patent 1143007, 1983.

5. Guibal, E.; Roulph, C.; Le Cloirec, P. Water Res. 1992, 26, 1139-45.

6. Macaskie, L.E.; Empson, R.M.; Cheetham, A.K.; Grey, C.P.; Skarnulis, A.J. Science1992, 257, 782-784.

7. Munroe, N.D.H.; Bonner, J.D.; Williams, R.; Pattison, K.F.; Norman, J.M.; Faison, B.D. In Abstracts, American Society for Microbiology Annual Meeting, 1993.

8. Hu, M.Z.-C.; Norman, J.M.; Faison, N.B.; Reeves, M. Biotechnol. Bioeng. 1996, 51, 237-47.

9. Horikoshi, T.; Nakajima, A.; Sakaguchi, T. Agric. Biol. Chem. 1979, 332, 617.

10. Byerley, J.J.; Scharer, J.M.; Charles, A.M. Chem. Eng. Journal 1987, 36, B49-B59.

11. Kuyucak, N.; Volesky, B. In Biosorption of Heavy Metals; Volesky, B., ed.; CRC Press: Boca Raton, FL, 1990, pp. 173-198.

12. Volesky, B.; Holan, Z.R. Biotechnol. Prog. 1995, 11, 235-250.

13. Kuyucak, N.; Volesky, B. Biorecovery 1989, 1, 189-204

14. Leusch, A.; Holan, Z.R.; Volesky, B. J. Chem. Tech. Biotechnol. 1995, 62, 279-288.

15. Aldor, I.; Fourest, E.; Volesky, B. Can. J. Chem. Eng. 1995, 73, 516-522.

16. Edgington, D.N.; Gorden, S.A.; Thommes, M.M.; Almodovar, L.R. Limnol. Ocean. 1970, 15, 945-955.

17. Fourest, E.; Volesky, B. Environ. Sci. Technol. 1996, 30, 277-282.

18. Crist, D.R.; Crist, R.H.; Martin, J.R.; Watson, J. In Metals-Microorganisms Relationships and Applications, FEMS Symposium Abstracts, Metz, France, May; Bauda, P., ed.; Societe Francaise de Microbiologie: Paris, France, 1993, p. 13.

19. Mullen, M.D.; Wolf, D.C.; Beveridge, T.J.; Bailey, G.W. "Sorption of heavy metals by soil fungi Aspergillus niger and Mucor Rouxii," In Soil Biol. Biochem. 1992, 24, 129-135.

Those predictive methods that are available in the open literature have limited or poor accuracy if applied to a wide variety of chemical systems and tower packings.{parse block="google_articles"}

The number of stages required for a given separation is obtained from the application of equilibrium thermodynamics. The actual number of stages obtained from a packed tower either in a laboratory, pilot plant, or an industrial plant is divided by the equilibrium stages predicted by vapor-liquid equilibrium thermodynamics to obtain an efficiency for the packed tower. Attempts have been made to generate semi-empirical correlations for packed tower efficiency from experimental data, and also generalized predictive models using the two-film theory of mass transfer. The mass transfer capability of a packing is typically expressed as HETP, HTU, K_Ga or K_La, all of which are rate-controlled quantities, and they can all be converted from one to another.

Attempts to derive generalized predictive methods for the mass transfer efficiency of packings using the two-film theory and dimensionless groups, and for the pressure drop and capacity using mechanistic models, have met with varying degrees of success. Published results of these attempts are the works of Bolles and Fair (1979), Bravo et al. (1987), Fair and Bravo (1987), Stichhnair et al. (1989), Fair and Bravo (1990), to name a few. The models used in these predictive methods were checked against many sources of pilot plant data, especially those made by Fractionation Research, Inc. (FRI) and the Separation Research Program (SRP) of the University of Texas at Austin.

On the other hand, reliable semi-empirical or empirical correlations of efficiency, capacity and pressure drop specific to a packing supplier's products can be found in their product bulletins, (e.g., Norton Chemical Process Products Corporation [NCPPC] 1987, 1992). These correlations are based on thermodynamic and physical properties of the systems, physical properties of the packings and numerous pilot plant tests and often operating data from industrial distillation columns. A very important need for ongoing pilot plant testing of tower packings in various distillation services arises because the existing predictive methods are either based on, or have been checked against only a limited data base i.e., limited number of chemical systems, system pressures (and temperatures) as well as packings. Thus pilot plant testing allows one to extend the database, which may suggest the need to refine the predictive models whether they are empirical, semi-theoretical or theoretical.

Often times, pilot plant distillation tests are necessitated because the customer requests such tests. The customer is anxious to have these tests performed because they want to minimum design and installation risk when building a multimillion-dollar facility. These risks can arise because of the lack of good vapor-liquid equilibrium data, the likelihood of azeotrope formation or interactions between key components not well understood, uncertainties in new design goats like high product purities even for familiar chemical systems, need to evaluate a new operating mode, etc.

The authors will discuss, based upon their experience in mass transfer tower design, operation of Norton's distillation pilot plants, and field feedback from the operation of commercial units, topics such as:

Packing size to tower diameter ratio
Distributor technology
Bed depth
Chemical system to be distilled
Sampling techniques
Reproducibility of results
Operation pitfalls

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Norton Distillation Pilot Plants

Norton Chemical Process Products Corporation (NCPPC) and its predecessor company has been operating a carbon steel distillation pilot plant for over 30 years at NCPPC's Chamberlain Laboratories in Stow, Ohio. The internal diameter of the tower is 387 mm (15.25 in.) and it could accommodate beds up to 3050 mm (10 ft.) in depth. {parse block="google_articles"}About seven years ago the height of this tower was raised so that it could accommodate a packed bed up to 6100 mm (20 ft.) in depth. This tower can operate at pressures in the range of 1.1 kPa (8 mm Hg. Abs.) to 170 kPa (10 psig). Most of the distillation data in the NCPPC data bank were collected using this carbon steel distillation column.

In 1992, NCPPC designed and built a new high-pressure distillation pilot plant. This tower and its ancillary equipment were fabricated from 316L stainless steel. This tower can be operated from high vacuum (0.133 kPa = 1 mm Hg. Abs.) to 2170 kPa (300 psig) at 177°C. It can be operated at pressures up to 2860 kPa (400 psig) at lower temperatures. This tower, like the carbon steel distillation tower has an internal diameter of 387 mm (15.25 in.). It can accommodate packed beds up to 7000 mm (23 ft.) in depth, resulting in a height-to-diameter ratio up to 18.

Figure 1: Distillation Pilot Plant

Both distillation pilot plants have similar flow schemes; they are located in a 12.2 m (40 ft.) tall high bay. The main difference is that the carbon steel distillation column sits atop a kettle reboiler, whereas the vapor produced in the stainless steel kettle reboiler enters the stainless steel distillation column through a 200 mm diameter side nozzle. A carbon steel skirt fastened to the floor supports the stainless steel column. Figure 1 shows the flow scheme of the stainless steel distillation tower, and Figure 2 is a scale drawing of the major pieces of equipment. Both columns can be operated at total reflux, or in the rectification mode at a LN ratio of less than 1. In addition, the high-pressure stainless steel tower has the capability to be modified as a center feed tower with beds up to 3050 mm (10 ft.) in the stripping as well as the rectification sections.

An important feature of the stainless steel tower is that it is provided with observation windows designed to withstand a pressure of 4236 kPa (600 psig) at 287°C. There are two pairs of windows in the vicinity of the reflux distributors and two pairs at the center feed location. One window of each pair is 100 mm in diameter used for illumination and the other window, which is perpendicular to the first, is 150 mm in diameter and is used for observation. The carbon steel distillation tower has three observation windows (150 mm diameter) in the vicinity of the reflux distributor. These observation windows permits the operator to observe the performance of the distributor, to look for any entrainment of liquid in the vapor and the onset of flooding. These windows have proved to be extremely valuable tools to characterize the distillation performance of the tower packings and distributors that have been tested over the years.

Figure 2: Elevation View of Pilot Plant

In the design of both pilot plants particular attention has been paid to minimize the hold-up of liquid in the overhead condensate circuit, viz., condenser, condensate line, condensate tank and reflux line. Both pilot plants use vertical condensers with the vapor condensing in vertical tubes thus minimizing the hold-up of the overhead product. The reboilers of both pilot plants are just large enough to hold sufficient charge of liquid such that the increasing hold-up of liquid in the packing resulting from increasing boil-up does not drastically deplete the reboiler liquid of its light component. The carbon steel reboiler can hold up to 0.38 cubic meters (100 gallons) of liquid and the stainless steel reboiler can hold up to 0.57 cubic meters (150 gallons) of liquid.

The members of FRI and SRP have access to the test data generated in the respective test columns. FRI has the capability to run high vacuum to high-pressure systems and the SRP can run systems from high vacuum to 507 kPa (60 psig), but neither FRI nor SRP has the capability to run corrosive systems. FRI has 1213 mm (47.75 in.) and 2438 mm (8 ft.) I.D. column sections whereas the SRP tower has an I.D. of 429 mm (16.875 in.). As far as the authors are aware of, NCPPC's stainless steel distillation pilot plant is the only one that is capable of testing all commercial size packings from high vacuum to high pressure in both non-corrosive and corrosive systems.

We have tested all sixes of NCPPC random packings in one of the two pilot distillation units. The tower diameter-to-packing size ratio ranged from 5.5 to 26. This list includes all sizes of Intalox Metal Tower packings (IMTP packing), Pall rings, Hy-Pak packing and several other random packings. Furthermore, all sizes of NCPPC's Intalox structured packings viz.; 1T, 2T, 3T, 4T and 5T, were also tested in these pilot distillation columns.

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Using Pilot Plant Data

The aims of most pilot distillation tests of a packing with a particular chemical system are to determine:{parse block="google_articles"}

• The mass transfer efficiency of the packing expressed as HETP or HTU
• The maximum hydraulic capacity and the maximum efficient capacity (MEC), i.e. the hydraulic capacity at which the efficiency starts to decline
• The pressure drop as a function of boil-up rate

Figure 3, which represents the typical HETP and AP vs. C_S, data for random packings and large structured packings are from Strigle and Rukovena, (1979). Figure 4 represents similar curves for small structured packings. Here C_S, is Souders and Brown (1934) entrainment parameter:

Eq. (1)

Where subscript "s" refers to Souders and Brown, G refers to "Gas", and L refers to "Liquid".

C_s = Entrainment parameters, m/s
G = Gas rate, kg/m² s
V = Gas velocity, m/s
ρ_G = Gas density, kg/m³
ρ_L = Liquid density, kg/m³

From the data of the type represented by Figure 3, the region B to C gives the design HETP and the point F gives the MEC. MEC represents the value of C, corresponding to the maximum rate at which the packing can be operated in distillation service while still maintaining the typical HETP as represented by the B to C portion of the HETP curve.

Figure 3: Random Packing Height Equivalent to a Theoretical Plate in Distillation Service, Strigle And Rukoneva (March, 1979)	Figure 4: Structured Packing Height Equivalent to a Theoretical Plate in Distillation Service

From the type of data as represented by Figure 4, i.e., for small structured packings, the MEC is determined by the C, value at which the slope of AP vs. C curve approaches infinity. The design HETP is taken at 90% of the MEC. Figures 5, 6, 7 and 8 show actual test data on No. 25 IMTP packing and No. 50 IMTP packing, Intalox Structured Packing 1T and Intalox Structured Packing 4T, respectively.

Figure 5: #25 Intalox Metal Tower Packing System System: Iso-octane / Toluene, 98.7 kPa Abs [740 mm Hg Abs] data by Norton	Figure 6: #50 Intalox Metal Tower Packing System System: Iso-octane / Toluene, 98.7 kPa Abs [740 mm Hg Abs] data by Norton

For most binary systems that we test in the pilot distillation columns, the number of stages generated in the packing is calculated using the method of Neretnieks et al. (1969). This method applies a coordinate transformation to the McCabe-Thiele method to account for the difference in the molal heats of vaporization between the two components. For multi-component systems, the stages are calculated using commercially available process simulators.

Figure 7: Intalox Structured Packing 1T System: Iso-octane / Toluene, 13.3 kPa Abs [100 mm Hg Abs] data by Norton	Figure 8: Intalox Structuted Packing 4T System: Iso-octane / Toluene, 13.3 kPa Abs [100 mm Hg Abs] data by Norton

The MEC point is confirmed by the heat balance on the distillation column. When the HETP starts to increase (decreasing efficiency) because of entrainment, this entrainment is carried into the condenser. This entrainment then manifests in the heat balance around the condenser as more heat being removed from the overhead based on the condensed vapor rate, than was put into the re-boiler.

The pressure drop across the packed bed is measured with the help of pressure taps above and below the packed bed and pressure transducers.

In operating an existing pilot plant distillation column, there are three fundamental issues involved:

a. Chemical test system
b. Tower packing
c. Liquid and or gas distributors

Typically, in a particular pilot test the performance data on two out of these three items are known fairly well; it is the purpose of the test to get information on the performance of the third item in the presence of the other two.

But in designing a new pilot distillation column one needs to decide ahead of time the type of chemical systems, whether corrosive or non-corrosive, high vacuum, atmospheric, or high pressure system that will be distilled in the column. As discussed earlier, the size and type of the packings and distributors to be tested will also have to be considered.

From an economic standpoint, the most important consideration is the diameter of the pilot distillation column.Download a legacy print-ready version of this article with high-resolution versions of Figures 1 and 2

Tower Diameter to Packing Size Ratio

Most of the early laboratory distillation data were taken in small columns, say 150 mm (6 in.) or less. Only very small random packings, viz., 3 to 12 mm (0.12- 0.5 in.) in size could be tested in such columns. There were several reasons for this. As the diameter of the pilot plant distillation column increases, in addition to the increase in installed cost, the cost of operating utilities, viz., reboiler steam and condenser cooling water increase proportional to the square of the tower diameter. Thus there is strong economic incentive for keeping the tower diameter as small as possible without affecting the quality of the test data.{parse block="google_articles"}

One consideration was the rule of thumb that the test tower diameter should be at least 10 times the size of the packing. The rationale behind this rule is that if larger packings were used, the wall area surrounding the packed bed would be a significant fraction of the packing area, and as such the column wall would provide a significant portion of the mass transfer area. It follows, based on this reasoning that when scaling up such data to large towers some derating would be necessary.

On the other hand, it can be argued that in a small tower, the gap between a bed of large packings and the tower wall can cause partial short-circuiting of liquid and vapor through these gaps. For structured packings, wall wipers minimize this problem. In large commercial towers the effect of such gaps will have negligible effect on packed bed hydraulics.

Most of the commercial size random packings fall in the size range of 15 mm to 90 mm (0.6-3.5 in.), and the structured packings has crimp heights in the range of 6 mm to 30 mm (0.25-l .2 in). But the majority of random packings sold commercially fall in the size range of 25 to 70 mm (1 - 2.8 in.), while the majority of commercially sold structured packings have crimp heights in the range of 8 mm to 12 mm (0.3-0.8 in.). With the 387 mm I.D. pilot distillation columns that NCPPC operates, it was found possible to test random packings in the size range of 15 mm to 70 mm (0.6-2.8 in.); the column I.D. to packing size ratio ranged from 26 to 5.5. In the case of the structured packings that were tested in these towers, the column I.D. to crimp height ratio ranged from 13 to 65. Based on experience with commercial installations, the test data taken in a 387 mm (15.25 in.) I.D. column gives reliable design data for commercial size columns. As mentioned earlier, the FRI columns have relatively large diameters, viz., 1213 mm (47.75 in.) and 2438 mm (96in.), probably because they were originally designed for testing trays. But the SRP columns which were built in 1986 have 429 mm (16.875 in.) I.D., because they were designed primarily for testing packings. Similarly another distillation pilot column operated by the Delft University of Technology in the Netherlands has an I.D. of 450 mm (17.72 in.) (Olujic et al., 1992). Thus, a distillation pilot column of approximately 400 mm (16 in.) I.D. gives reliable test data on random and structured packings. This type of test data along with reliable distribution technology can be used, without any scale-up factor, to design commercial distillation columns.

Distribution Technology

Factors to be considered in selecting liquid distributors for a distillation test tower are:

• Turndown ratio and height of the distributor
• The number and size of distribution points (orifices) per unit tower cross-sectional area
• The liquid flow variation allowed between distribution points
• The layout of the liquid distributor points over the tower cross-sectional area

It is common practice, when testing a packing, to cover the complete operating range of the packing. In the authors' experience, the typical turndown ratio is 5:l. And, it is not uncommon to have a 7:l turndown ratio. Several types of liquid distributors are used for distillation tests. Except for the notched weir-trough distributor (which happens to have high turn-down ratio), spray distributor (which is seldom used in distillation), most of the distillation distributors fall into one of the following three categories.

• Orifice-pipe arm distributors
• Orifice-pan distributors
• Orifice-trough distributors

Let us first consider the design of orifice-plate and orifice-trough distributors. Both of these types of distributors are open at the top. In the orifice pan distributor, the gas flows through specially designed risers as well as the area between the pan and the tower wall. The rest of the pan area is available for locating liquid orifices. In the orifice-trough distributor, the liquid is held in specially designed troughs with liquid orifices at the bottom and/or on the sides of the troughs; the rest of the tower cross-sectional is available for gas flow.

For a given orifice size, the flow rate through the orifice is approximately proportional to the square root of the liquid head, when the orifice is running full of liquid. Therefore, for a given set of orifices at a fixed elevation, the required head of liquid above the orifices is proportional to the square of the liquid flow rate. Thus a 2: 1 turndown ratio in flow requires a 4: 1 ratio of liquid head. Typically the minimum liquid head required for predictable flow of liquid through the orifice is about 50 mm (2 in.). Thus the liquid head required at maximum flow rate for 2: 1 turndown is 200 mm (8 in.). For 5: 1 turndown the maximum required is 1250 mm (50 in.), and for 7: 1 turndown the maximum head required is 2450 mm (8 ft.). It follows that, unless over 2.5 m (8 ft.) of column height can be reserved for liquid distributor, one must resort to using a distributor with multiple levels of orifices or use more than one single-level orifice distributor, each with a different orifice size. The design features of many of these types of distributors are proprietary.

The pipe-arm distributors depend, for their performance, on the liquid head prevailing upstream of the orifices; this pressure is generated usually by a liquid feed pump. The turndown capability of the pipe-arm distributors are only limited by the capacity of the feed pump and the maximum allowable velocity of liquid through the orifices above which formation of liquid spray might cause entrainment. The biggest drawback of this type of distributor is that the flow variation from orifice to orifice can be excessive, especially at high flow rates due to variability of the size and shape of the orifices and the pressure drop through the pipe arms. Therefore, orifice-pan and orifice-through distributors are generally preferred for both pilot plant distillation columns and industrial distillation columns.

The number of liquid distribution points required for unit tower cross-sectional area is a function of the type and size of the packing. Based on the authors' experience, the following general statements can be made:

• Large random packings require fewer pour points than smaller random packings.
• Large structured packings require fewer pour points than medium sized structured packings.
• Small structured packings have better liquid spreading characteristics than larger structured packings.
• Except for small random packings, most packings will operate well with pour point densities of between 40 points/m² (4/ft²) and 60 points/m² (6/ft²). Even small random packings of commercial interest perform well with 100 points/m².
• The smallest size orifice used is 2-3 mm in diameter; this small orifice can only be used with clean systems.
• Sufficient liquid head should be allowed to limit the individual orifice flow variation to ± 5% of the average flow rate.
• The layout of liquid distributor orifices over the tower cross-sectional area is based on the method of Moore and Rukovena (1986)

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Bed Depth

Several considerations go into choosing the maximum allowable depth of a packed bed. {parse block="google_articles"}Maximum number of theoretical stages generated in a bed e.g., 15 theoretical stages per bed is a rule of thumb used often. But as Table I shows we have observed packed beds of Intalox Structured Packing 1T generating over 21 theoretical stages in a single bed, irrespective of the pour point density.
Table 1 also shows that a single bed of Intalox Structured Packing 2T can generate as many as 32 stages in a single bed.

Height-to-diameter ratio of the bed again, as can be seen from Table 1, the authors have observed that a bed of Intalox structured packing can operate well with a height-to-diameter ratio of up to 15.

Another consideration is the mechanical strength of the packing and the support system for a tall bed. A packed bed depth chosen based on the two criteria listed above can be supported without any problem.

Table 1: Intalox Structured Packing Performance

Chemical Systems to be Distilled

A large number of pilot plant tests are performed during the various stages in the development of a new packing. Since the primary purpose of these tests is to compare the performance of the new packing with other packings, all tests are performed with one or a few chemical systems. For example, Zuiderweg (Circa 1966) lists a number of test mixtures that can be used at atmospheric, vacuum and pressure distillation. FBI typically uses the cyclohexane-heptane system at 34.5 kPa (260 mm Hg Abs), atmospheric pressure and 165 kPa (24 psia), the p-xylene/o-xylene system at 13.3 kPa (100 mm Hg Abs) and lower, and the i-butane/n-butane system at high pressures ranging from 689 kPa (100 psia) to 2758 kPa (400 psia). The SRP typically uses the test system cyclohexane/n-heptane at 33.3 kPa (250mm Hg Abs), atmospheric pressure, 165 kPa (24 psia) and 413 kPa (60 psia).

Historically, NCPPC has used the iso-octane/toluene system at atmospheric pressure and 13.3 kPa (100 mm Hg Abs) as the standard test system. We also have used the cyclohexane-heptane system and p-xylene/o-xylene system as standard test systems.

In the carbon steel tower, NCPPC has tested our packings with numerous other chemical systems, e.g., methanol/water, cyclohexanone/cyclohexanol, acetone/water, water/MEG, to name a few. We also have tested numerous proprietary systems for our customers over the years. Using the data generated from our test columns, as well as commercial installations, NCPPC has developed, efficiency, capacity and pressure drop correlations, which have been used to design world-class high vacuum to high-pressure distillation columns using NCPPC packings.

In the future NCPPC will be adding, to its data bank, test data on high pressure and corrosive systems taken in our high-pressure stainless steel distillation column.

Sampling Techniques

Generation of HETP data from distillation tests requires drawing samples from the system after steady state has been attained. For example, during a distillation test in the rectification mode, liquid samples are drawn from the overhead vapor condensate, from below the packing and from the reboiler. The overhead sample is collected from the discharge side of the reflux pump. The sample from under the packing is collected using a trough type sampler in the shape of a cross; the liquid leaves the sampler from the center and is conducted through tubing to the outside. These samples are chilled, if necessary, to avoid loss due to vaporization and consequent change in composition. At a given boil-up rate, the onset of steady state is monitored by drawing samples periodically, say every l/2 hour, and analyzing the samples. For most binary organic systems, a gas chromatograph is the most convenient analytical tool. The difference between the compositions of the light component in the overhead sample minus that in the packing sample increases gradually until it reaches an asymptotic value at steady state. During the run, three consecutive samples are taken at half hour intervals. Typically, for a good run at steady state HETP values calculated using the three samples differ from one another by no more than 8 mm (0.3 in.). For example, in Figure 5, for every run (i.e., C_s) three HETP values and delta P measurements were obtained. It can be seen that for the majority of runs the three measurements coincide with one another. The time required to attain steady state, after changing the boil-up rate, is typically about two hours for common organic systems with relative volatilities above -1.1. But systems with relative volatility approaching 1, e.g. isotopes, it can take from 16 to 24 hours for the attainment of steady state composition profile in the packed bed. The reboiler sample, together with the sample drawn from under the packing, is used to calculate the number of stages generated in the reboiler - which is usually around 1; this procedure is used as a check on the accuracy of the sample drawn from under the packed bed.

Reproducibility of Test Results

Factors that affect the reproducibility of test results for a given packing, without considering the manufacturing tolerance of the packing are numerous.

Method of packing the distillation column. For structured packings, care should be taken to see that the wall wipers properly engage the tower wall so that bypassing of vapor and liquid through the gap between the packing and the tower wall are minimized, if not eliminated. It is important that consecutive layers of structured packing are rotated by a fixed angle, usually 70" with respect to each other, so that the seams of segmental bundles do not line up (for a 387 mm (15.25 in.) I.D. test tower each layer is made as a single piece). We make sure that no gap is allowed to exist between the packing and manways. A plug contoured in the shape of the tower wall is pressed against the packing. We have noticed that, in the absence of this plug, short-circuiting of liquid and vapor can result in poor packing efficiency.

Analysis of samples - In a typical distillation test in which gas chromatography is used for analyzing the samples, it is important that standards are run every day. It is necessary to make a sufficient number of standards to cover the anticipated range of composition bracketing that of the overhead sample and the re-boiler sample. In a binary mixture, for example, the response factors of the two components can vary over the range composition. The response factors can be drastically different when the composition approaches pure light component and pure heavy component compared to those of a 50-50 mixture.

Insulation of the tower walls- In both our distillation pilot plant, (150 mm (6 in.) thick fiberglass blanket insulation with aluminum wrap is used to cover the re-boiler and the tower wall to minimize condensation of the internal vapor traffic at the tower wall; this condensation would otherwise affect the internal reflux ratio. We try to limit the heat loss to about one percent of the heat input to the reboiler.

Accurate and reproducible pressure drop measurement requires careful design of pressure taps and lines leading to the pressure transducers. It is important to make sure that any vapor that condenses in the lines flows back to the tower without affecting the pressure measurement. The pressure taps are designed so that the opening faces downwards to prevent liquid from entering the tap. A baffle is provided in the opening to prevent vapor from impinging on the opening. This baffle ensures that only static pressure is measured. Typically 12 mm (0.5 in.) diameter tubing, which continuously pitches from the transducer to the pressure tap, assures that any vapor condensing in the line runs back to the tower.

The authors have found that it is possible to obtain reproducible HETP data on the same system-packing combination, after repacking the tower and recharging the system, with variation not exceeding 15 mm (0.6 in.).Download a legacy print-ready version of this article with high-resolution versions of Figures 1 and 2

References

1. Bolles, W.L. and Fair, J.R., I Chem. E symposium series No. 56, p. 3.3135, 1979
2. Bravo, J.L., Rocha, J.A. and Fair, J.R., Hydrocarbon Processing 65, P. 45, March 1986{parse block="google_articles"}
3. Fair, J.R. and Bravo, J.L., Chemical Engineering Progress 86, p. 19, 1990
4. Fair, J.R. and Bravo, J.L., I. Chem. E. symposium Series No. 104, p. Al83, 1987
5. Matthews, M., Hukill Chemical Corporation, Bedford, Ohio, Private Communication, 1990
6. Moore, F. and Rukovena, F., "Liquid and Gas Distribution in Commercial Packed Towers", 36th Canadian Chemical EngineeringConference", Paper 236, October, 1986.
7. Neretnieks, I., Ericson, I. and Eriksson, S., British Chemical Engineering 14, 12, p. 653, 1969
8. Norton Chemical Process Products Corporation (NCPPC), "Intalox High-Performance Structured Packing", Bulletin, 1992
9. Norton Chemical Process Products Corporation (NCPPC), "Intalox High-Performance Systems", Bulletin IHP-1, 1987
10. Olujic, Z., Stoter, F. and De Graauw, J., AIChE First Separation Division Topical Conference on Separation Technologies, New Developments and Opportunities, November 2-6, Miami Beach, Florida
11. Souders, M. and Brown, G.G., Industrial and Engineering Chemistry 2, p. 98, 1934
12. Stichlmair, J., Bravo, J.L. and Fair, J.R., Gas Separation Purification 1, p. 19, 1989
13. Strigle, R.F. and Rukovena, F., Chemical Engineering Progress 75, No. 3, p. 86, 1979
14. Zuiderweg, F.J., "Recommended Test Mixtures for Distillation Columns", European Federation of Chemical Engineering, Working Party on Distillation, Absorption and Extraction, Circa 1966.

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