CHAPTER 2 REVIEW OF LITERATURE -...

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CHAPTER 2

REVIEW OF LITERATURE

2.1 GENERAL

The use of solar energy for the production of food, fibre and heat

has been known to mankind for a long time. Research over the last five

decades has also made it possible to produce mechanical and electrical power

with solar energy. Although the potential of solar radiation for disinfection

and environmental mitigation has been known for years, only last two

decades, has this technology been scientifically recognized and researched.

When sunlight is used to cause a chemical reaction by direct absorption, the

process is called photolysis. If the objective is achieved by the use of

catalysts, it is known as photocatalysis. Photocatalysis by titanium dioxide has

been demonstrated to be an inexpensive and effective method for treating a

variety of organic pollutants in water (Augugliaro et al 2004). The UV

radiation required for photocatalytic processes may come from an artificial

source or the sun.

The artificial generation of UV radiation contributes to a large

portion of the operating, capital and maintenance costs of a photocatalytic

reaction system because of the utility consumption and periodic replacement

of the UV lamps. There is, therefore, a significant economic incentive to

develop solar powered photocatalytic reactors. In addition, the environmental

impact induced by the use of solar energy is minimal and this renders the

photocatalytic process environmentally attractive (Chan et al 2003). The

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application of solar powered photocatalytic reactors to treat water

contaminated with organic pollutants holds promise for regions receiving

strong sunlight throughout the year, like India.

Advanced Oxidation Processes (AOPs) involve high cost if compared

to biological treatment processes. However, the use of AOPs is more suitable

when the wastewater to be treated is not easily biodegradable. One

economically viable option to treat wastewater containing non-biodegradable

pollutants consists of combining an AOP, for instance photocatalysis and a

biological post treatment. In this case, the photocatalysis step is used to

enhance the biodegradability of the wastewater, so that it can be more easily

treated biologically. Photocatalysis has been suggested to be feasible and

promising to treat wastewaters containing non-biodegradable wastewaters,

being used as a pre-treatment method to increase the biodegradability (Sarria

et al 2003).

2.2 ADVANCED OXIDATION PROCESSES (AOPs)

Advanced Oxidation Processes (AOPs) generally involve

generation and use of powerful but relatively non-selective transient oxidizing

species, primarily the hydroxyl radical (˙OH), in some vapour-phase

advanced oxidation processes, singlet oxygen has also been identified as the

dominant oxidizing species. Table 2.1 shows that ˙OH has the highest

thermodynamic oxidation potential, hence ˙OH based oxidation processes

have gained the attention of many advanced oxidation technology developers.

In addition as shown in Table 2.2, most environmental contaminants react

1 million to 1 billion times faster with ˙OH than with O3, a conventional

oxidant (US EPA 1998).

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Table 2.1 Oxidation potential of several oxidants in water

Sl.No. Oxidant Oxidation Potential (eV)

1 Hydroxyl radical ( .OH) 2.80

2 Singlet oxygen O(1D) 2.42

3 Ozone (O3) 2.07

4 Hydrogen peroxide (H2O2) 1.77

5 Perhydroxy radical 1.70

6 Permanganate ion 1.67

7 Chlorine dioxide (ClO2) 1.50

8 Chlorine (Cl2) 1.36

9 Oxygen (O2) 1.23

Table 2.2 Rate constants for O3 and ˙OH Reactions with organic

compounds in water

Sl.No. Compound Type Rate Constant (M-1 s-1)

O3 ˙OH

1 Acetylenes 50 108 to 109

2 Alcohols 10-2 to 1 108 to 109

3 Aldehydes 10 109

4 Alkanes 10-2 106 to 109

5 Aromatics 1 to 102 108 to 1010

6 Carboxylic acids 10-3 to 10-2 107 to 109

7 Chlorinated alkenes 10-1 to 103 109 to 1011

8 Ketones 1 109 to 1010

9 Nitrogen-containing organics 10 to 102 108 to 1010

10 Olefins 1 to 450 103 109 to 1011

11 Phenols 103 109 to 1010

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2.2.1 Classification of Advanced Oxidation Processes

The AOP may be classified depending on the source to generate the

oxidizing species. The commonly used classification is shown in Figure 2.1.

The Advanced Oxidation Processes are Photolysis, AOPs based on ozone,

AOPs based on hydrogen peroxide, Hot AOPs, Photocatalysis, electro-

chemical oxidation, Ultrasound technologies and electron beam oxidation

(Peratitus et al 2004, and Gogate and Aniruddha 2004 a). This research work

mainly focused on photocatalysis using solar light as irradiation source.

Figure 2.1 Classifications of advanced oxidation processes

Advanced Oxidation Processes

- Photolysis - Ozone based AOPs - H2O2 based AOPs - Hot AOPs - Photocatalysis - Ultrasound technologies - Electro-chemical oxidation - Electron beam oxidation

- Photolysis UV - Photolysis VUV - Ozonation at alkali conditions - Ozone + UV and/or H2O2 - Ozone + Catalyst - Fenton’s Reagent - Fenton-like Reagent - Photo-Fenton Reagent - H2O2 / UV Reagent - Electro-Fenton - Supercritical Wet Oxidation - Wet Oxidation - Wet Peroxide Oxidation

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2.2.2 Photocatalysis

Among AOPs, heterogeneous photocatalysis has proved to be of

real interest as efficient tool for degrading both aquatic and atmospheric

organic pollutants. Heterogeneous photocatalysis involves the acceleration of

photoreaction in presence of semiconductor photocatalyst. One of the major

applications of heterogeneous catalysis is photocatalytic oxidation (PCO) to

effect partial or total mineralisation of gas phase or liquid phase pollutants to

benign substances (Gaya and Abdullah 2008).

2.2.3 Mechanism of photocatalysis

Basically, if a photon has more energy than the band gap of a

material, it will free up an electron when absorbed. The band gap of a material

represents the difference in the energy of the electrons in the valence band of

the atom and the conduction band. As an electron moves up from the valence

band to the conduction band and becomes free, it leaves a positive hole

behind. The positive hole and the negative electron may recombine with the

release of thermal energy, unless they interact with neighbouring atoms of

other materials to cause chemical reactions. Such reactions are known as

photoreactions, since they are initiated by photons. If the reaction involves

atoms or molecules that act as catalysts, the reaction is known as

photocatalytic (Gaya and Abdullah 2008, and Bhatkhande et al 2001).

Sunlight may be used in both photolytic and photocatalytic reactions that

could result in useful applications, such as the oxidation of toxic organic

chemicals or production of hydrogen. The energy of a photon is given by

the Equation (2.1).

hch

(2.1)

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where h is the Plank’s constant (6.625 10 -34 J-s), is the frequency, λ the

wavelength and c is the speed of light (3 108 m/s).

Titanium Dioxide (TiO2) has a band gap of 3.2eV, therefore the

wavelength λ of a photon with energy equal to the band gap of TiO2 is given

by the Equation (2.2).

hc

(2.2)

= 34 86.625 10 3 10

3.2

= 0.388 10 -6 m or 388 nm

Therefore, a photon of sunlight with a wavelength of 388 nm or

less (i.e. energy 3.2 eV or higher) will excite an electron from a valence band

(vb) to a conduction band (cb) when absorbed, resulting in a free electron (e-)

and a positive hole (hole+) as per Equation (2.3).

TiO2 + h (photon) → holevb+(TiO2) + e-

cb(TiO2) (2.3)

The holes (hole+) and the electrons (e-) are both highly energetic

and mobile. They may recombine and release heat or migrate to the surface.

On the surface, they may react with adsorbed molecules of other species and

cause a reduction or oxidation of that species. Since recombination, in the

bulk or near the surface, is the most common reaction, the quantum yields

(molecules reacted/photons absorbed) of most photolytic reactions are low.

Separation of the electron-hole pairs is aided by formation of a potential

gradient near the surface of the semiconductor. This space charge region

results from the different electrical potential of the solid semiconductor and

the liquid phase of the ambient solution. For TiO2 this potential drives valance

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band holes toward the particle surface and conduction band electrons away

from the surface. The electrons and holes at the surface become active sites

for oxidation and reduction of adsorbed molecules, as illustrated in Figure 2.2.

Figure 2.2 Mechanism of photocatalysis

The positive holes cause oxidation of the surface adsorbed species

while the electrons cause reduction. Both reactions must take place in order to

maintain electro neutrality. Thus, if the objective is the oxidation of organics,

the electrons must be consumed in a reduction reaction such as absorption by

oxygen molecules to form a superoxide, in order to keep the holes available

for oxidation. On the other hand, if the objective is the reduction and recovery

of metals, all other reducible species, such as oxygen, must be eliminated or

kept away. The following Equations (2.4) and (2.5) describe the oxidation and

reduction reactions.

TiO2 + h → TiO2(e-cb+hole+

vb) (2.4)

TiO2(e-cb+hole+

vb) TiO2 + (heat) (2.5)

recomb

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If the objective is oxidation of an organic pollutant, the electron is

consumed by an adsorbed compound/molecule (Equation 2.6).

O2 + e-cb ═> O2

-˙ (2.6)

If a water molecule is adsorbed on the surface of TiO2, it may exist

as hydroxyl and hydrogen ions (Equation 2.7).

H2O ═> OH- + H+ (2.7)

The negatively charged hydroxyl ion adsorbed on the surface of

TiO2 gives up its negative charge (electron) to the positive hole to regenerate

the neutral TiO2 and in turn becomes a neutral hydroxyl radical (Equations

(2.8) and (2.9)).

TiO2 (hole+vb) + OH-

ads ═> OH˙ + TiO2 (2.8)

2O2-.+2H+

aq → H2O2 + O2 (2.9)

H2O2 may yield additional hydroxyl radicals by any of the

following reactions (Equations (2.10) - (2.12))

H2O2 + e-cb→ OH˙ + OH- (2.10)

H2O2 + O2-. → OH˙ + OH- + O2 (2.11)

H2O2 +h → 2OH˙ (2.12)

The hydroxyl radical is a very potent oxidizing agent which can

oxidize a pollutant organic molecule ClHmXn, into CO2 and H2O, directly or

through intermediate compounds as per Equation (2.13).

ClHmXn +y OH˙ +zO2 → lCO2 + mH2O + nHX (2.13)

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Table 1.1 shows the oxidation power of various species relative to

that of chlorine. The oxidation power of the hydroxyl radical is ranked second

among these known strong oxidizing agents. This demonstrates the hydroxyl

radical’s potential for oxidizing normally hard to destroy pollutants like

halogenated organics, surfactants, herbicides and pesticides (Gaya and

Abdullah 2008).

2.2.4 Factors Influencing Photocatalysis

Several chemical and physical parameters have been shown to

affect the photocatalytic process. The main chemical parameters include pH,

initial pollutant concentration, catalyst dosage and oxygen while physical

parameters include temperature and light intensity. Many works have been

done to investigate the effects of such parameters on the photocatalytic

kinetics (Bhatkhande et al 2001, Gogate and Aniruddha 2004, and Gaya and

Abdullah 2008).

2.2.4.1 Effect of direct photolysis

Some pollutants can only be dissociated in the presence of UV

light. For this, the pollutant must absorb the light of the lamp (or the sun) with

a reasonable photodissociation quantum yield. Although organic pollutants

absorb light over a wide range of wavelengths, this is generally stronger at the

lower wavelengths. However, such natural photodegradation is usually very

slow. For example, at least 10 days under perfectly sunny conditions is

necessary to reduce 50 mg/L acrinathrin to half. One half of 100 mg/L

pentachlorophenol at pH 7.3 decomposes in 48 hours. So the photolytic

reaction rate is usually different from one substance to another even in the

same experimental device. These tests have to be performed in order to find

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out the decomposition rates without the semiconductor. This type of

experiment will focus on demonstrating the absence or evaluating the

importance of the following effects viz., the treatment is not feasible without a

catalyst, increase in temperature (due to illumination) in the reactor does not

cause product loss, there is no adsorption of pollutant or its metabolites in the

materials of the pilot plant.

2.2.4.2 Effect of oxygen

In semiconductor photocatalysis for water treatment, the pollutants

are usually organic and, therefore, the overall process can be summarised by

Equation (2.14). Given the reaction stoichiometry of this equation, there is no

photomineralization unless O2 is present. It is necessary for complete

mineralization and does not seem to be competitive with other reactives

during the adsorption on TiO2 since the places where oxidation takes place are

different from those of reduction. The O2 avoids the recombination of e-/h+

and, photoactivated oxygen (O2•-) also reacts directly (Table 1.1). The

concentration of oxygen also affects the reaction rate, which is faster when

the partial pressure of oxygen (pO2) in the atmosphere in contact with the

water increases (Gogate and Aniruddha 2004 a).

Organic pollutant + O2 → CO2 + H2O +mineral acids (2.14)

2.2.4.3 Effect of pH

The pH of the aqueous solution significantly affects TiO2, including

the charge of the particle and the size of the aggregates it forms. The pH at

which the surface of an oxide is uncharged is defined as the Zero Point

Charge (pHzpc), which for TiO2 is around 7. Above and below this value, the

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catalyst is negatively or positively charged according to Equations (2.15) and

(2.16).

-TiO2H2+ ↔ TiOH + H+ (2.15)

-TiOH ↔ TiO- + H+ (2.16)

In consequence the photocatalytic degradation of organic

compounds is affected by the pH (Tsong and Wan 1991).

2.2.4.4 Effect of catalyst dosage

Either in static or in slurry or dynamic flow photoreactors, the

initial reaction rates was found to be directly proportional to catalyst dosage.

This indicates a truly heterogeneous catalytic regime. However, above a

certain value, the reaction rate levels off and becomes independent of catalyst

dosage. This limit depends on the geometry and working conditions of the

photoreactor and is for a definite amount of TiO2 in which all the particles,

i.e. the entire surface exposed, are totally illuminated. When catalyst dosage is

very high, after travelling a certain distance on an optical path, turbidity

impedes further penetration of light in the reactor. In any given application,

this optimum catalyst dosage has to be found in order to avoid excess catalyst

and ensure total absorption of efficient photons (Velegraki and Dionissions

2008).

2.2.4.5 Effect of temperature

Because of photonic activation, photocatalytic systems do not

require heating and operate at room temperature. The true activation energy Et

is nil, whereas the apparent activation energy Ea is often very low (a few

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kJ/mol) in the medium temperature range (20º-80°C). However, at very low

temperatures (-40°- 0°C), activity decreases and activation energy Ea becomes

positive. By contrast, at high temperatures (>70-80°C) for various types of

photocatalytic reactions, the activity decreases and the apparent activation

energy becomes negative. When temperature increases above 80°C, nearing

the boiling point of water, the exothermic adsorption of reactants is

disfavoured and this tends to become the rate-limiting step (Gaya and

Abdullah 2008). In addition to these mechanical effects, other consequences

of plant engineering must be considered. If temperature is high, the materials

used for the plant should be temperature-resistant (more expensive), and

oxygen concentration in water decreases. Consequently, the optimum

temperature is generally between 20 and 80°C. This absence of need for

heating is attractive for photocatalytic reactions carried out in aqueous media

and in particular for environmental purposes (photocatalytic water

purification).

2.2.4.6 Effect of light intensity

At low light intensity and corresponding low carrier concentrations,

the rate of oxidation of a particular compound is proportional to light

intensity, while at higher light intensity the rate is dominated by second-order

charge carrier combination and has a square-root dependence on light

intensity. The transition from one regime to the other depends on the

photocatalyst material, but is typically above 1 sun equivalent. This transition

depends on the catalyst configuration and on the flow regime in the

photoreactor, and varies with each application. The optimal light power

utilization corresponds to the domain where the destruction rate is

proportional to light intensity (Chen and Ray 2001).

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2.2.4.7 Effect of initial pollutant concentration and kinetic model

In heterogeneous catalysis, the pollutants have to be adsorbed on

the catalyst surface sites for bond breaking or formation. The adsorption of

pollutants and the availability of sites are hence important parameters in

photocatalytic reactions. The rate of pollutants conversions is proportional to

the active sites. As the reaction proceeds, the amount of pollutants adsorbed

on the catalysts surface will decrease until the pollutants is completely

converted (Velegraki and Dionissios 2008).

Kinetic models are often formulated to describe photocatalytic

reactions with respect to the initial pollutant concentrations. The kinetic

models for photocatalytic reactions are derived based on the classical

heterogeneous catalysis model, which is the Langmuir-Hinshel wood (LH)

kinetic model. This model assumes that the reaction occurs on the surface and

the reaction rate (r) is proportional to the fraction of surface coverage by the

pollutant is given by Equation (2.17).

r tdc Kcr k kdt 1 Kc

(2.17)

where kr is the reaction rate constant, K is the adsorption constant and C is the

pollutant concentration at any time t. Integrating the above Equation (2.17)

yields

rColn K (Co C) k KtC

(2.18)

where Co is the initial pollutant concentration. At high Co, the Equation

(2.18) can be simplified to a zero order rate law (Equation 2.19)

r(Co C) k Kt kt (2.19)

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At low Co, a first order rate Equation (2.20) can be obtained from

Equation (2.17)

rColn k Kt ktC

(2.20)

where k is the apparent rate constant.

2.2.4.8 Effect of anions and cations

Industrial wastewater contains apart from the pollutants, different

salts at different concentrations. The salts are generally ionized under the

conditions of photocatalytic degradation (Robert et al 1999 and Alhakimi et al

2003 b). The anionic and cationic parts of the salts have different effects on

the photocatalytic degradation process. The presence of anions such as

chlorides, sulphate, carbonate and bicarbonate is common in industrial

wastewater. These ions affect the adsorption of the degrading species, act as

hydroxyl ion scavengers and may also absorb UV light (Kamble et al 2004).

2.3 OVERVIEW OF WOR KDONE ON PHOTOCATALYSIS

Photocatalysis, i.e. using semiconductor particles under bandgap

irradiation as little microreactors for simultaneous reduction and oxidation of

different redox systems has been intensively studied during the last 30 years,

since the pioneering work of Carey et al in 1976. The photocatalytic

degradation processes are gaining importance in the area of wastewater

treatment, because these processes result in complete mineralization with

operation at mild conditions of temperature and pressure (Gogate and

Aniruddha 2004 b). A major advantage of the photocatalytic oxidation based

processes is the possibility to effectively use sunlight or near UV light

(Kanmani et al 2003). Many researchers (Bhatkhande et al 2003, Marcpera

et al 2004, Moon et al 2005, Linda et al 2005, Kaniou et al 2005, Yuan et al

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2006, Priya and Giridhar 2006, Muruganandham et al 2006 a, and Tamer et al

2007) have shown that most of the organic pollutants like 4-cholorophenol,

nitrobenzene, alachlor, Lanasol blue dye, sulfamethazine, microcystin-LR,

phenol, reactive yellow and p-nitrophenol are completely oxidizated into non

toxic products of carbon dioxide and water. Various catalysts such as TiO2,

ZnO, CdS, ZnS, etc have been used as photocatalyst so far in different studies

reported in the literature (Siemon et al 2002, Sarria et al 2003, Marcpera et al

2004, Gogate and Aniruddha et al 2004, Linda et al 2005).

The surface area and the number of active sites offered by the

catalyst for adsorption of pollutants plays an important role in deciding the

overall rates of degradation as usually the adsorption step is the rate

controlling. It should be noted that the best photocatalytic performances with

maximum quantum yields have been always with titanium dioxide. In

addition, Degussa P25 catalyst is the most active form among the various

ones available and generally gives better degradation efficiency (Chan et al

2001 and Linda et al 2005). The energy needed to activate TiO2 is 3.2 eV or

more, which corresponds to near UV radiation of a wavelength of 388 nm or

less. As 4-6 percent of sunlight reaching the earth’s surface is characterized

by these wavelengths, the sun can be used as the illumination source.

The stages involved in the development of photocatalysis assume

cyclic form. The cycle, shown in Figure 2.3 starts and ends with the analysis

of needs by the end-user. Each stage of the development cycle is critical for

successful operation of the next immediate step. Some of the illustrative work

done in the recent years in the area of photocatalysis applied to wastewater

treatment has been furnished in Table 2.3 with discussion about reactor

operating conditions and the important findings obtained in the work. It was

observed that, most of the Photocatalytic Oxidation (PCO) studies reported in

the literature, used single-constituent model solution with a concentration up

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to 100 mg/L. But, the concentration of phenolic compounds in industrial

wastewaters varies from 100 to 500 mg/L and also contains inorganic

pollutants.

Figure 2.3 Photocatalytic system development cycle

Majority of academic studies reported, have used artificial UV light

source for photocatalytic oxidation. Solar light intensity in our region is quite

suitable and can be used as a light source in the PCO process. By employing

solar light, the common drawback of relatively high cost of UV lamps and

electricity can be overcome. Hence it is necessary to study, the PCO process

using solar energy and prove it to be an economically and technologically

feasible option for degradation of non-biodegradable organic pollutants in

industrial wastewaters. Though, advanced oxidation processes oxidize almost

all pollutants, considering the economical aspects, the use of advanced

oxidation process alone as treatment may not look lucrative. Thus, a hybrid

method consisting of using advanced oxidation process to reduce the toxicity

of the wastewaters up to a desired level followed by biological treatment is

perhaps needed for the future.

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Table 2.3 Typical findings observed in the representative works related to photocatalytic degradation of wastewaters

Sl. No.

Compounds / Wastewater

Treated

Experimental condition and Type of equipment Important findings of the work Reference

1. Acid Green 16 (910-4 M)

Experiments were conducted in an immersion type photochemical reactor equipped with 8 W low pressure mercury lamp.

At 5 10-4 M of dye the degradation efficiency was 97.5 % whereas for 10 x10-4 M the efficiency was 79.45 %. The photodegradation efficiency increases rapidly with increasing amounts of TiO2 from 100 mg to 250 mg and then decreases on further increase of TiO2. It was observed that photodegradation efficiency was more in acidic pH. A solution containing 9 10-4 M of the dye could be completely decolourised in 3 min and 90 % degradation could be achieved in 3 h.

Sakthivel et al (2000)

2. Fiber manufacturer wastewater (COD=114-134 mg/L)

A Pyrex cylindrical reactor with total volume of 3.5 L equipped with 365 nm UV lamp was used.

The experimental results indicated that with the TiO2 as catalyst, the optimal condition was 0.25 g/L. Both increased light intensity and continuous aeration increased COD removal efficiency, particularly under continuous aeration for significantly raising the ratio of BOD/COD to improve efficiency of subsequent biological treatment.

Hsieh et al (2000)

3. Azo dye (10 mg/L)

Experiments were conducted in a 1 L pyrex glass batch reactor equipped with 4 nos. of UV lamps.

A Pseudo first order reaction kinetic was proposed to simulate the photocatalytic degradation of orange G in the batch reactor. More than 80% of 10 mg/L orange G decomposition in 60 minutes reaction time was observed in this study and fast decomposition of orange G only occurred in the presence of both TiO2 and suitable light energy. Faster degradation of orange G was achieved under acid conditions. The degradation rates of orange G at pH 3 were about two times faster than those at pH 7. The reaction rates were proportional to TiO2 concentration and light intensity with the power order of 0.726 and 0.734 respectively.

Hung et al (2001)

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Table 2.3 (Continued)

Sl. No.


Treated


4. EDTA (5.0 mM)

Irradiations were performed using a high-pressure xenon arc lamp.

Experiments with 5.0 mM EDTA were performed for 3 hours irradiation under different conditions. Under N2 bubbling, depletion of EDTA was very low. Under O2 bubbling, the concentration of EDTA decreased around 90 %. However, the corresponding decrease of TOC ranged only between 4.5 % and 9 %.

Babay et al (2001)

5. Humic acid (TOC = 32 mg/L)

Irradiation experiments were carried out with 9 numbers of 15 W low-pressure UV lamps.

It was confirmed that the technique used in this study was effective to remove TOC at 38 % and colour at 89 % within 150 min oxidation. The experiment results showed that low concentrations of hydrogen peroxide dosage (less than 0.016 M) to UV/ TiO2 system accelerated the TOC and colour removal rate from 9 % to 38 % and 40 % to 89 % respectively, while over dosage made this positive effective decline.

Tay et al (2001)

6. Methyl orange (2x10-3-1x10-5 M)

0.5 L three necks round bottom flasks were used as reactors. solar radiation was used as an irradiation source.

It was found that 0.4 % of TiO2 gave the highest degradation rate constant (0.619 h-1). In the second set of experiments TiO2 concentration was fixed at 0.4 % and the Methyl Orange was varied, the highest rate constants was obtained with the concentration of Methyl Orange was 4x10-5 M and it was found to be 0.639 h-1. The degradation becomes negligible in the presence of high concentrations of Methyl Orange. The highest degradation rate was obtained at pH =3 with a rate constant k = 2.6683 h-1 followed by that at pH=9 with a rate constant k = 0.7585 h-1.

Alqaradawi and Salman (2002)

7.

Humic acids (100 mg/L)

Experiments were carried out in a solar box. The light source was Xenon lamp.

The adsorption study of Humic acids at three pH Solution (1.9/7.5/11) indicated that at acidic pH, Humic acids are adsorbed on TiO2. It was obtained 88 % of TOC removal after 6 h of irradiation with optimum TiO2 loading 1.0 g/L. It was shown that Photocatalysis process improved the biodegradability of Humic acids.

Wiszniowski et al (2002)

31


Sl. No.


Treated


8. Landfill leachate (COD =1260-1673 mg/L)

Quartz photoreactor 190 mL was equipped with 8 W Mercury lamp.

Under the acidic condition, the photocatalytic removal efficiency of landfill leachate was relatively high because there was very low concentration of inorganic carbons that could inhibit the photocatalytic oxidation. The effect of TiO2 concentration was obtained from experimental results and it could demonstrate the relationship of amounts of TiO2 dosage and reaction rate.

Cho et al (2002)

9. Benzidine and 1,2-dipehnylhydrazine (1 mM)

Irradiation experiments were conducted in a lab scale reactor of 250 ml capacity equipped with 500 W high-pressure mercury lamp.

The maximum efficiency for the degradation and the TOC depletion was observed at pH 5. The efficiency for the degradation as well as for the mineralization of the compound increases with the increase in the substrate concentration from 0.1 – 1 mM. The catalyst concentration was investigated using different concentrations varying from 0.5 to 5 g/l. The degradation and mineralization rate increases with the increase in catalyst concentration from 0.5 to 2 g/l. Bromate ions found to enhance the rate of degradation.

Muneer et al (2002)

10. Potassium hydrogen phthalate (200 ppm)

The photo reactor used for the degradation process was a 3L beaker made of Pyrex glass, equipped with a magnetic stirrer and an oxygen-purging device.

The optimum catalyst concentration was found at 3 g/L of Degussa P25 and 10 g/l of Hombikat UV 100. The optimum degradation rate was obtained at pH 5.0 for both Catalysts. The complete degradation of organic pollutants achieved after 5 h of irradiation using the optimized reaction conditions.

Alhakimi et al (2003 b)

11 Phenol and substituted phenols.

A Pyrex reactor open to air equipped with 125 W lamp.

Direct UV-irradiation of TiO2 catalyst suspended in aqueous solutins is able to destroy hydroxyphenol and nitrophenols. The photodegradation of these phenolic compounds follows Pseudo-first-order kinetics. It was found that the degradation of nitrophenols could be accelerated in acidic medium whereas hydroxy phenols are less sensitive to pH variations. It was also found that the BOD5 in the phenolic solutions increased with the COD decreased during photodegradation reaction.

Ksibi et al (2003)

32


Sl. No.


Treated


12 4-Chlorphenol (100 ppm)

Photoreactor consisted of a 3L beaker made of Pyrex glass, equipped with a magnetic stirring bar and an oxygen purging device.

The comparison of the degradation of 4-chlorophenol, using sunlight and a UV lamp was carried out using two different catalysts, Degussa P25 and Hombikat UV 100. The optimum concentrations for Degussa P25 and Hombikat UV 100 occurred at 7 and 10 g/L, respectively. The optimum initial pH was found to be 5 for both catalysts. The degradation rate of 4-chlorophenol is 6.4 times and 1.6 times higher when using sunlight compared to the artificial UV lamp for Degussa p25 and Hombikat UV 100, respectively. The degradation rate of 4-chlorophenol is six times higher, compared to Hombikat UV 100 at the optimum conditions, when using sunlight and Degussa P25 as the catalyst.

Alhakimi et al (2003)

13 Inactivation of E. coli

Pyrex glass bottle of 50 mL was used as batch reactor. Irradiation by simulated sun test lamp.

Parameters such as light intensity, extend of continuous irradiation, catalyst concentration and temperature have a positive effect on disinfection. Intermittent illumination results in an increase in the time required for E. coli inactivation. No illumination of a contaminated TiO2 suspension. In contrast, without catalyst, illuminated bacteria recovered it initial concentration after 3 h in the dark. Bacterial inactivation in the absence of Catalyst was more affected than that with catalyst. When increasing light intensity from 400 to 1000 W/m2. TiO2 concentrations higher than 1 g/L do not increase the initial inactivation rate for both intensities. Water turbidity negatively affects the photocatalytic inactivation of bacteria.

Rincon and Pulgain (2003)

33


Sl. No.


Treated


14 Sulfonylurea herbicides

Irradiation experiments were carried out with 125 W mercury medium pressure UV lamps, in a Pyrex Cylindrical flask opened to air.

Results show that adsorption is an important parameter controlling the apparent kinetic order of the degradation. In respect to the Langmuir – Hinshelwood model, the photocatalytic reaction is favoured by a high concentration. At moderate light fluxes (13.3 mW/cm2), the rate is proportional of photonic flux () giving optimum quantum efficiency. For higher values (39.8 mW/cm2), the electron – hole recombination becomes predominant decreasing the quantum efficiency although the reaction rate still increases.

Vulliet et al (2003)

15 Imazaquin (herbicide) (50x10-6 mol/ L)

Irradiation experiments were carried out with 125 W Mercury lamp in a 70 ml cylindrical borosilicate glass reactor.

A lower solution pH in the 3-11 range was found to be favourable to degradation. The addition of H2O2 up to 10-3 mol / L enhanced the degradation rate and decreased it at higher concentrations. The photocatalytic effect was more efficient in a suspension containing 2.0 g/L TiO2 with 1 h sonication in the dark, rather than with 20 min. sonication before irradiation. Solar radiation decomposed the herbicide faster than an artificial light source.

Garcia and Keiko (2003)

16 Biocides (2 mg/L)

Lab-scale photoreactor consisted of 1500 W Xenon arc lamp.

The primary degradation of the micro pollutants follows a Pseudo-first-order kinetics following the Langmuir-Hinshelwood model. The total disappearance of chlorothalonil and dichlofluanid was achieved in 90 and 20 minutes respectively, whereas the materialization of organic carbon to carbondioxide after 240 minutes of irradiation was found to be 100 % for chlorothalonil and 78 % for dichlofluanid.

Sakkas and Albanis (2003)

17 Dicamba (2,5-Dichloro-6-methoxybezoic acid)

Experiments were conducted in a quartz reactor of 500 mL capacity and illuminated with 150 W UV lamps.

Photolysis reactions were slow but the corresponding photocatalysis rates were increased by about 3 and 5 times in the presence of TiO2 at 300 and 350 nm UV, respectively. Photocatalytic rates were increased with the pH at acidic to neutral ranges. The results of H2O2 assisted photocatalysis experiments showed that a low H2O2 dosage in photocatalysis would enhance the decay rate of dicamba by 2.4 times, but an overdose of H2O2 will retard the rate.

Chu and Wong (2004)

34


Sl. No.


Treated


18 Azo dyes (20 mg/L)

Experiments were conducted in a 3 L cylindrical glass reactor equipped with a 8W low pressure Mercury lamp.

The study examined degradation of azo dyes using photocatalytic oxidation. The photocatalyst used were TiO2, ZnO and SnO2. The reaction rate constants fits a first order reaction model and the reaction rate constant for TiO2 + SnO2 (0.31 h-1) is larger than that of TiO2 (0.24 h-1). The reaction rate constants had higher values at pH 10 than pH 7. At pH 10, the trend of the reaction rate constants of azo dyes followed the order: ZnO > ZnO + SnO2 > TiO2 + ZnO > TiO2 + SnO2 > SnO2.

Wu (2004)

19 Reactive azo dye (50 mg/L)

Irradiation experiments were carried out in an open batch reactor of 1000 mL capacity equipped with an 300 W UV lamp.

The results obtained from the experiments adding H2O2/ TiO2 show that the highest decolorisation rate is provided by the combination of (UV+ TiO2 +H2O2). The decolorisation efficiencies were 18 %, 22 %, 34 % and 52 % in the runs UV, UV+H2O2, UV+ TiO2 and UV+ TiO2 + H2O2

after approximately 100 min irradiation respectively. The decolourisation rate constant was 0.018 min-1 in the 1 g/L TiO2 while it was 0.004 min-1 in the presence of 0.125 g/L TiO2. The results of the obtained oxygen uptake rate measurements in biological activated sludge have shown that the photocatalytically treated dye was easier to degrade than untreated dye.

Nevim (2004)

20 Bisphenol-A (10 mg/L)

Experiments were performed in an immersion type reactor equipped with a 6 W lamp. The reactor capacity was 1.5 L.

The effects of immobilized TiO2 –film thickness, UV radiation intensity and pH on the photodegradation were investigated. Apparent rate constant of the first order increased with increasing TiO2 coating time from 1 to 3, however, decreased over 4-coating time. Rate constant increased as the pH value shifted from basic to acidic regions.

Lee et al (2004)

21 Protozoan, fungal and bacterial microbes

Solar irradiation was simulated using a 1000 W Xenon arc Solar simulator.

The ability of solar disinfection (SODIS) and solar photocatalytic disinfection (SPC-DIS) batch-process reactors to inactivate water borne protozoan, fungal and bacterial microbes waste evaluated. After 8 h simulated solar exposure both SPC-DIS and SODIS achieved at least a 4-log unit reduction in viability against Protozoa, fungi and bacteria. The inactivation is approximately 50 % faster in batch process SPC-DIS reactors than SODIS reactors under comparable conditions.

Lonnen et al (2005)

35


Sl. No.


Treated


22 Alachlor (herbicides) (2 mg/L)

Experiments were conducted in an immersion type reactor of 1L capacity equipped with 400 W metal halide lamp.

The photodegradation rate of alachlor could be described as an apparent first order. The rate constant of alachlor increased from 0.021 to 0.060 h-1 as the number of coating times increased from 1 to 5 times in the absence of ferric ion, where the corresponding thickness of TiO2 film were 67 and 174 nm. The rate constant increased from 0.030 to 0.060 h-1 as pH value 9 to 5. The rate constant increased slightly from 0.031 to 0.050 h-1 as the concentrated of ferric ion increased from 0.75 to 7.5 mg/L in the absence of TiO2. However those increased from 0.051 to 0.110 h-1 in the presence of both TiO2 and ferric ion. In situ photodeposition of ferric ion onto the TiO2 surface enhanced the rate constant of photodegradation of alachor by about 80 % with an adding 7.5 mg/L.

Kim et al (2005)

23 Reactive Black (3.85 10 -4 M)

Open borosilicate glass tube of 50 mL capacity was used as reactor. solar as light source

The rate of degradation was maximum at catalyst amount 2 g/L, pH 9. A complete removal of 3.85 10 -4 M dye solution under solar irradiation was observed in 3.5 h. It was observed that after 90 min. the removal efficiency was 70 and 57 % for photocatalytic and photochemical process.

Muruganan-dham et al (2006 b)

24 p-toluenesulfonic acid (100 mg/L)

Experiments were carried out in a quartz cylindrical reactor of 100 mL Irradiation by solar.

The rate of degradation was maximum at catalyst amount 1 g/L, pH 3 and concentration 50 mg/L. It was also observed that in the presence of anions and cations, the rate was decreased.

Kamble et al (2006)

36


Sl. No.


Treated


25 Reactive yellow (5 10-4 mol/L)

Multilamp photoreactor of capacity 50 mL with 8 W lamp was used.

The rate of degradation was maximum at catalyst amount 4 g/L, pH 5.5 and dye concentration 5x 10-4 mol/L. The increase of radiation intensity from 16 to 62 W increases the decolourisation from 32.9 to 87 % in 20 min. The presence of electron acceptors like H2O2, KBrO3 enhanced the degradation. Negative effects are observed in the presence of NaCl and Na2CO3.

Muruganan-dham et al (2006 a)

26 Textile dye house wastewater (COD = 404 mg/L)

Immersion batch reactor of capacity 1 L with 400 W high pressure Hg lamp was used.

At pH 3 and catalyst amount 0.5 g/L, complete removal of colour and 40-60 % COD after 4 h of treatment was observed. It was observed that efficiency of catalyst marginally deteriorated on repeated use. H2O2 enhanced the reaction rate was observed.

Pekakis et al (2006)

27 Dye solutions (50 mg/L)

Irradiation experiments were carried out in a 1.5 L UV reactor with 125 W high pressure Hg lamp.

The rate of degradation was maximum at catalyst amount 1g/L, pH 7 and concentration 50 mg/L. Complete removal of dye was achieved with in 100 min.

Bizani et al (2006)

28 Basic dye (50 mg/L)

Irradiation experiments were carried out in a 1 L reactor with 12 nos. of 8 W lamps.

The results showed that agitation speed varying from 50 to 200 rpm has an influence on the dye decomposition rate. The decomposition rate increases with TiO2 amount up to 0.98 g/L and then decreases. And also it was observed that the rate increases with the UV power intensity up to 64 W and then decreases.

Wu et al (2006)

29 Tannery wastewater (COD = 800 mg/L)

Annular batch reactor of 2 L capacity irradiated by 80 W high pressure Hg lamp was used.

The application of AOP (H2O2/UV, TiO2/H2O2/UV and TiO2/UV) was investigated. The COD removal increased in the order UV < H2O2/UV < TiO2/H2O2/UV < TiO2/UV treatments. In H2O2/UV treatment, the COD removal reached around 60 % in 4 h reaction.

Sauer et al (2006)

30 Estrone and 17 β estradiol (500 mg/L)

Batch reactor of 400 mL capacity with 150 W high pressure UV lamp.

The maximum removal of 97 % was observed at pH 7.6, catalyst amount 1 g/L and pollutant concentration 500 mg/L with in 4 h of irradiation. The presence of humic acid increases the degradation rate.

Zhang et al (2007)

37


Sl. No.


Treated Experimental condition and

Type of equipment Important findings of the work Reference

31 Olive mill wastewater (1300 mg/L)

The reactor had a working volume of 100 cm3. It composed of water cooled double-walled Pyrex glass tube. It was illuminated by a vertical UV tube of 415W.

Olive mill wastewater (OMW) was treated by photocatalysis using TiO2 under UV irradiation on the laboratory scale. The photocatalyst was TiO2 Degussa P25. It was added at a level of 1 g/L. The chemical oxygen demand, the coloration at 330 nm, and the level of phenols all showed decreases which, after a 24-h treatment, reached 22 %, 57 %and 94 %, respectively.

Hajjouji et al (2008)

32 Benzoic acid (150 mg/ L)

Experiments were carried out in a cylindrical reaction vessel, batch type, coupled with a double-walled immersion well. A medium pressure mercury lamp providing 25mW/cm2 was used as light source.

Experiments were conducted at benzoic acid initial concentrations between 25 and 150 mg/ L, catalyst loadings between 0.2 and 1 g/ L and initial solution pH values between 2 and 10.6. Conversion increased with increasing catalyst loading up to about 0.6 g/L and it was favoured at alkaline or neutral conditions but impeded at extremely acidic conditions. Increasing initial substrate concentration led to decreased benzoic acid conversion, which was found to follow a Langmuir–Hinshelwood kinetic expression. Complete conversion can be achieved in less than about 60 min.

Velegraki and Dionissios (2008)

33 Black table olive wastewater (COD = 1-8 g/L)

Experiments were conducted in an immersion well, batch type, laboratory scale photoreactor, equipped with a 400W high pressure mercury lamp.

Initial organic load, expressed in terms of chemical oxygen demand (COD), was studied in the range 1–8 g/L, anatase TiO2concentrations in the range 0.25–2 g/L and H2O2 concentrations in the range 0.025–0.15 g/L. Treatment efficiency, which was assessed in terms of COD, total phenols, aromatics and colour reduction, generally increased with decreasing initial COD and increasing contact time, catalyst and H2O2 concentrations. Depending on the conditions employed, nearly complete decolouration (>90%) could be achieved, while mineralization never exceeded 50 %.

Chatzisymeon et al (2008)

34 Winery wastewater (800 mg/L)

Experiments were carried out in batch operation in an annular type reactor. The lamp used was a medium pressure mercury arc UV lamp of 435 nm.

The performance of the reactor was studied as a functional of various operating variables, such as gas flow rate, pH and catalyst loading. It was found that the optimum gas flow rate was 6 L/min whereas the optimum pH value is 6.5. The highest photodegradation rate and the maximum COD removal were achieved at zero catalyst loading with COD removal of about 84 %. Lower rates of chemical reaction in photocatalysis compared to photolysis were possibly because of the shielding of UV light by catalyst particles.

Agustina et al (2008)

38

2.4 SOLAR PHOTOCATALYTIC REACTORS

The artificial generation of photons required for the treatment of

wastewater is the most important source of costs during the operating of

photocatalytic wastewater treatment plants. This suggests using the sun as an

economically and ecologically sensible light source with a typical UV-flux

near the surface of the earth of 20 - 30 W/m2 the sun puts 0.2 - 0.3 mol

photons/m2/h in the 300 - 400 nm range at the process disposal (Bahnemann

2004). Principally these photons are suitable for destroying water pollutants in

photocatalytic reactors. Over the last two decades several reactors for the

solar photocatalytic water treatment have been developed and tested.

Designs of solar photocatalytic reactors have followed the well-

known designs of solar thermal collectors including concentrating and non

concentrating designs. The key differences are viz., the fluid to be treated in

the reactors must be exposed to UV solar radiation, therefore, the absorber

must be transparent to UV solar radiation; and no insulation is needed, since

temperature does not play a significant role in the photoreaction. As a result,

the first engineering scale outdoor reactor developed was a simple conversion

of a parabolic trough solar thermal collector. The conversion replaced

absorber/glazing tube combination of the thermal collector with simple pyrex

glass tube through which contaminated water can flow. This reactor was used

to treat water contaminated with Trichloroethylene (TCE). The catalyst, TiO2

powder was mixed with contaminated water to form slurry which was passed

through the pyrex glass tube (reactor tube) located at the focal line of the

parabolic trough. Since that time, a number of reactor concepts and designs

have been advanced by researchers all over the world (Sagawe et al 2003 a

and Sagawe et al 2003 b). Based on the method of collecting sunlight, two

reactors systems are designed viz., concentrating and Non-concentrating

39

reactor systems. Based on the condition of the catalyst, two reactor systems

are designed and they are suspended and fixed catalyst systems.

2.4.1 Concentrating Collectors

Solar photochemical processes are based on the collection of only

high-energy short-wavelength photons to promote photochemical reactions.

Most of the solar photochemical processes use UV or near-UV sunlight (300-

400 nm), but in some photochemical synthesis processes, upto 500 nm

sunlight can be absorbed and photo-Fenton heterogeneous photocatalysis uses

sunlight up to 580 nm. Sun light at wavelengths over 600 nm is normally not

useful in any photochemical process. Nevertheless, the specific hardware

needed for solar photochemical applications has much in common with those

used for thermal applications. As a result, both photochemical systems and

reactors have followed conventional solar thermal collector designs, such as

parabolic troughs and non-concentrating collectors.

The original solar photoreactor designs for photochemical

applications were based on line-focus parabolic-trough concentrators (PTCs).

In part, this was a logical extension of the historical emphasis on trough units

for solar thermal applications. Furthermore, PTC technology was relatively

mature and existing hardware could be easily modified for photochemical

processes. The first outdoors engineering – scale reactor developed (in USA)

was a converted solar thermal parabolic-trough collector in which the

absorber/glazing-tube combination had been replaced by a simple pyrex glass

tube through which contaminated water could flow. The first engineering-

scale solar photochemical facility for water detoxification in Europe was

developed by CIEMAT using 12 two-axis PTCs, each consisting of a turret

and a platform supporting four parallel PTCs, with an absorber at the focus of

each collector. The platform has two motors controlled by a two-axis

40

(azimuth and elevation) tracking system. Thus, the collector aperture plane is

always perpendicular to the solar rays, which are reflected by the parabola

onto the reactor tube (concentrating ratio ≈ 10) at the focus through which

circulates the contaminated water to be detoxified (Curco et al 1996, Malato

et al 1996).

Typical overall optical efficiencies obtained in this PTC were

around 50 %. Parabolic-trough collectors make efficient use of direct solar

radiation and, as an additional advantage, the thermal energy collected from

the concentrated radiation could simultaneously be used for other

applications. The reactor is small, while receiving a large amount of energy

per unit volume. The flow is turbulent and volatile compounds do not

evaporate, so that handling and control of the liquid to be treated is simple

and cheap. The main disadvantages are that the collectors (i) use only direct

radiation, (ii) are expensive, and (iii) have low optical and quantum

efficiencies. Several different substances have been successfully degraded

with these collectors viz. chromium (VI), dichloroacetic acid, phenol,

4-chlorophenol (CP), dichlorophenol (DCP), pentachlorophenol (PCP),

atrazine, and industrial wastewaters (Malato et al 2004).

2.4.2 Non-Concentrating Collectors

Non-concentrating collectors are, in principle, cheaper than PTCs

as they have no moving parts or solar tracking devices. They do not

concentrate radiation, so that efficiency is not reduced by factors associated

with concentration and solar tracking. Manufacturing costs are cheaper

because their components are simpler, which also means an easy and low-cost

maintenance. Also, the non-concentrating collector support structures are

easier and cheaper to install and the surface required for their installation is

smaller, because, since they are static, they do not project shadows on the

41

others. An extensive effort in the design of small non-tracking collectors has

resulted in the testing of several different non-concentrating solar reactors

viz., plate reactor, tubular reactor, falling film reactor, and shallow solar pond

reactor (Goswami et al 1997).

Although non-concentrating collectors are simple, the design of a

robust non-concentrating photoreactor is not trivial, since they must be

weather-resistant, chemically inert and UV transmissive. In addition, flow in

non-concentrating systems is usually laminar, which presents mass transfer

problems and vaporization of reactants. Consequently, the use of tubular

photoreactors has a decisive advantage because of the inherent structural

efficiency of tubes, which are also available in a large variety of materials and

sizes and are a natural choice for a pressurized fluid system.

2.4.2.1 Flat-plate reactors

This reactor consists of a back plate on which water trickles down

from a spray bar at the top. A woven mesh (fibreglass) covers the back plate

in order to damp out surface waves as the water trickles down and to even out

the flow. This allows a thin even film of water. A UV transparent glazing

prevents any evaporation from the flow. This design allows contaminated

water, mixed with the catalyst particles as slurry, to be treated as the water

trickles down when exposed to the sun. The reactor can be operated in a fixed

catalyst configuration by replacing the plain woven mesh with a mesh with

the catalyst fixed on it (Wyness et al 1994 a).

2.4.2.2 Tubular reactors

A simple reactor concept consists of transparent inflatable tubes

connected in parallel between two headers. As the water flows through the

42

reactors under pressure, the tubes inflate. Large areas of this reactor can be

rolled or folded into a small volume for portability and transported for on-site

use. Such tubular reactors were tested at the University of Florida to treat

water contaminated by volatile organic compounds. These reactors were also

tested in the field at Tyndall Air Force Base to treat groundwater

contaminated with fuel, oil and lubricants (Goswami et al 1997).

2.4.2.3 Shallow pond reactor

Shallow pond-type reactors, developed at the University of Florida,

may be constructed on-site especially for industrial waste water treatment.

Since industries already use holding ponds for microbiological treatment of

wastewater, shallow solar ponds can be used for the front end or the back end

of a combined solar/microbiological treatment of wastewater. These reactors

would be ideal for wastewater treatment in industries such as pulp and paper,

textiles, pharmaceuticals, and chemicals (Wyness et al 1994 b). The reactor

can be operated in slurry or fixed catalyst configuration. If TiO2 is used as a

catalyst in a slurry configuration, it settles down to the bottom, if it is not

continuously mixed. While the disadvantage of this configuration is that

continuous mechanical mixing is needed, the advantage is that after the

catalyst settles down, the treated water can be removed from the top without

filtration. Wyness et al (1994 a) tested shallow pond reactors in both slurry

and fixed configuration, and various area-to-depth ratios, for treating

contaminated water. The reactors worked extremely well under various

insulation conditions (sunny, partly cloudy and cloudy).

2.4.2.4 Falling-film reactor

The vertical film of this reactor is open to the atmosphere on both

sides, which allows it to make maximum use of the diffuse atmospheric

43

radiation. The integrity of the film is maintained by means of vertical strings

appropriately spaced apart (Goswami 1997).

2.4.2.5 Thin film fixed bed reactor

One of the first solar reactors not applying a light concentrating

system and thus being able to utilize the diffuse as well as the direct portion

of the solar UV-A irradiation for the photocatalytic process is the thin film

fixed bed reactor. The most important part of the thin film fixed bed reactor is

a sloping plate of width 0.6 m height 1.2 m coated with the photocatalyst

(TiO2) and rinsed with the polluted water in a very thin film (~ 100 µm). The

flow rate is controlled by a cassette peristaltic pump and can be varied

between 1 and 6.5 L/h (Bockelmaan et al 1993).

2.4.2.6 Double Skin Sheet Reactor (DSSR)

It is a non-concentrating reactor and it consists of a flat UV

transparent structured box made of polymethyl metha acrylate (plexiglass). It

can utilise both the direct and diffused solar radiations. One module has a

length of 1.4 m, a width of 0.98 m and contains 30 channels each of which is

28.5 mm 12 mm . The total inner volume of one module is 14.4 L. In all

experiments , this system was operated in a recirculation mode. The turbulent

flow rate was adjusted to 11.8 L resulting in a Reynolds number of about

9000. The residence time in the reactor was 73 sec. The reactor was adjusted

with an inclination angle of 45˚ (Goslich et al 1997 and Bahnemann et al

1999).

44

2.4.2.7 Coated Mesh Reactor (CMR)

It consists of a freshly coated 6.5 m long by 0.5 m wide TiO2

impregnated woven glassfibre mesh supported on an inclined (20˚) corrugated

support forms the main element of the coated mesh reactor. The coated mat

had a TiO2 density of 5 g/m2 and the reactor was operated in a continuous

recirculatory mode. A 500 L capacity reservoir tank was located at the base of

the reactor, which supplied 200 L of solution to the inlet of the reactor (Feitz

et al 2000).

2.4.2.8 Packed Bed Reactor (PBR)

It was constructed from mirror finished stainless steel to provide

maximal internal reflectivity and is approximately 1 2 m2 in aperture area.

The reactor vessel contains 6.5 cm deep TiO2 coated rashing rings and

perforated baffles to direct the flow and maximize contact time. Air is

bubbled through reactor section to ensure sufficient mixing and oxygen

exposure (Feitz et al 2000).

2.4.2.9 Multiple Step Cascades Falling Reactor (STEP)

It was composed of 21 stainless steel stairs, stair height 70 mm,

stair width 500 mm, covered with a 1 m2 pyrex sheet to limit water

evaporation. The effluent flowed over the steps before being collected in a

tank, from which it was elevated with a pump to the top of the steps for

recirculation (Guillard et al 2003).

45

2.4.2.10 Water bell fountain reactor

It was developed by constructing a water bell fountain, which is

transparent to solar light in a shallow pond. The water fountain was created by

means of a fountain nozzle assembly installed at the centre of the reactor.

Through an opening provided at the bottom of the reactor, the wastewater

containing the photocatalyst was withdrawn from the reactor and pumped to

the central tube of the fountain nozzle by using a circulation pump (Kanmani

et al 2005).

2.4.3 Compound Parabolic Concentrator (CPC)

Compound parabolic concentrators extensively employed for

evacuated tubes, are an interesting cross between trough concentrators and

one – sun systems and are a good option for solar photochemical applications.

CPCs are static collectors with an involute reflective surface around a

cylindrical reactor tube. They have been found to provide the best optics for

low concentration systems and can be designed with a concentrating ratio

close to one, thus having the advantages of both PTCs and one-sun collectors,

since these static collectors can capture both direct and diffuse UV- sunlight.

The advantages of the solar CPC system is its intrinsic simplicity, while

it is also cost-effective, easy to use, and requiring low capital investment. The

reflector design enables almost all the UV – radiation arriving at the CPC

aperture (not only direct, but also diffuse) to be collected and available to the

process in the reactor. The light reflected by the CPC is distributed around the

back of the tubular photoreactor so that most of the reactor tube

circumference is illuminated. Because of the CPC aperture-to-tube diameter

ratio, the incident light on the reactor is very similar to that of a one-sun

photoreactor. As in a parabolic trough, water is easier to pipe in and to

46

distribute than in many one-sun designs. All these factors contribute to

excellent CPC collector performance in solar photochemical and

photocatalytic applications (Bahnemann et al 1999).

2.5 OVERVIEW OF WORK DONE IN THE DEVELOPMENT

OF PHOTOCATALYTIC REACTORS

To ensure efficient conversion of incident photons to charge

carriers, the appropriate design of solar photocatalytic reactor is most

important. Some laboratory and pilot plant photoreactor configuration have

been proposed for photocatalysis. The most typical laboratory-scale

configurations include vigorously stirred batch photochemical reactor,

cylindrical flasks, glass dishes and annular differential photoreactors. Based

on the condition of photocatalyst, two reactor systems are designed viz.

suspended and fixed catalyst systems. Some of the illustrative work done in

the development of photocatalytic reactors applied to wastewater treatment

has been furnished in Table 2.4 with discussion about reactor operating

conditions and the important findings obtained in the work. Matthews (1986)

reported the use of suspended TiO2 in a solar illuminated tubular reactor

equipped with a parabolic trough concentrator to degrade a variety of organic

pollutants. The light concentrating reactor requires reflectors, which are

expensive and only able to concentrate the direct component of the solar

irradiation for reaction. This leads to the conclusion that scaling up such a

design may be difficult.

The first engineering scale solar photochemical facility for water

detoxification in Europe was developed by CIEMAT using

12 two axis PTC. But the main disadvantages are that collectors use only

direct radiation, expensive and have low optical and quantum efficiencies.

Several different substances have been successfully degraded with these

47

collectors viz. dichloroacetic acid, phenol, 4-chlorophenol, dichlorophenol

and atrazine. In contrast, non-light concentrating reactors utilizing both the

direct and diffuse components have greater potential for process development.

Non-concentrating collectors are cheaper than PTC as they have no moving

parts or solar tracking devices. Wyness (1994 a) developed Flat plate reactor

and tested the reactor in an out-door solar photocatalytic oxidation facility.

Goswami (1994) reported that all the three non- concentrating reactors viz.

Flat plate, Shallow pond and Tubular reactors demonstrated satisfactory

performance in solar photocatalytic oxidation facilities when tested over a

wide range of operating conditions. Goslich et al (1997) developed Double

Skin Sheet Reactor (DSSR) and Thin Film Fixed Bed Reactor (TFFBR). But

it requires large catalyst area for purification of wastewater and also

constrained by mass transfer limitations due to laminar flow conditions. Feitz

et al (2000) developed fixed catalyst reactors viz. coated mesh and packed

bed reactor and its efficiency was very less compared to suspended catalyst.

Kanmani et al (2003) developed two water fountain solar

photocatalytic reactors of suspended catalyst system for treating textile dyeing

wastewaters. Chan et al (2003) constructed a solar photocatalytic cascade

reactor to study the photocatalytic oxidation of benzoic acid. Though a lot of

studies have been reported so far, still the efficient use of reactors at large-

scale are lacking due to opacity, light scattering and depth of radiation

penetration. Engineering design and operation strategies are lacking for

efficient use of reactors at large scale. Moreover, the requirement of at least

one side to be transparent to UV light significantly poses size limitations

along with breakage risk. Hence, studies are needed in terms of design of

simple reactor achieving uniform irradiation and minimal losses of incident

light due to opacity, light scattering and adsorption by liquid.

48

Table 2.4 Typical findings observed in the representative works related to solar photocatalytic reactors

Sl. No. Types of Reactor

Pollutant/ Wastewaters

Description of Reactor Important findings Reference

1. Coated mesh reactor (CMR)

Phenol (2 mg/L)

A freshly coated 6.5 m x 0.5 m TiO2 impregnated woven glass fibre mesh supported on an inclined corrugated support forms the main element of the reactor. The reactor operated in a continuous recirculatory mode. The flow rate was 5 L/min and irradiation was by solar light

The total removal of phenol from the solution for an exposure period of 200 min was only 36 %. Although the overall processing rate is slow, the removal rate over the length of the reactor surface was rapid indicating the deficiency of the rector arrangement, i.e., long non-contact periods due to the small reactor to tank volume ratio.

Feitz et al (2000)

2. Packed-bed reactor (PBR)

Phenol (2mg/L) The reactor vessel was constructed from mirror finish stainless steel and approximately 1x2m2 in aperture area. The reactor contains 6.5 cm deep TiO2 – coated Raschig rings and perforated baffles. The flow rate was 3 L/min and irradiation was by solar light.

PBR achieved 98 % removal of phenol from a 2 mg/L concentration, 100 L solution after 3 h irradiation. The air injection system forms an important component of PBR treatment. Under clear sky conditions with no air injection, 75 % of phenol was removed over 4 h compared to 99.3 % removal with additional air input.

Feitz et al (2000)

49


Sl. No.

Types of Reactor



3. Solar Thin film Cascade rector

Benzoic acid (25-100 mg/L)

Nine stainless steel flat places were arranged in a Cascade configuration. The plates were coated with TiO2 catalyst. The plates were positioned 2.5 cm vertically apart from each other and inclined at an angle of 5˚ to the horizontal. The area of each plate exposed to solar light was 17.5 cm 28 cm.

The water fall effect introduced by the cascade design can promote mass transport and aeration in the solution film. The percentage removal of TOC in 7 L of 100 mg/L benzoic acid solutions increased from 30 % to 83 % by adding 10 mL of hydrogen peroxide solution. The average TOC removal rates did not demonstrate significant dependence on TOC0 and the intensity of the light was found to be the dominant factor affecting the degradation process.

Chan et al (2003)

4. Thin film cascade rector (UV)

Benzoic acid (100 mg/L)

Photoreactor consists of six fluorescent UV lamps and a cascade of three flat plates coated with TiO2 catalysts. The plates were positioned 2.5 cm vertically apart from each other and inclined at an angle of 5˚ to the horizontal. The area of each plate exposed to solar light was 17.5 cm 28 cm

The percentage removal of TOC for the flow rates from 2 L/min to 5 L/min was 7.34, 8.15, 7.14 and 7.95 respectively, which seemed to show no dependence of the removal of TOC on the flow rate. Increasing temperature from 291K to 300 K was found to decrease the DO level in the solution, hence reducing the percentage removal of TOC. When the number of lamps decreases from six to two the UV intensities decreases from 1.39 mW/cm2 to 0.56 mW/cm2 and the removal of TOC also decreases from 19.43 % to 11 %. The removal of TOC slightly decreased from 19.43 % for the cascade reactor to17.1 % for the single plate reactor.

Chan et al (2001)

50


Sl. No.

Types of Reactor



5. Solar immobilized-catalyst photocatalytic reactor

Microcystin-LR (100 mg/L)

The reactor vessel was constructed from mirror-finish stainless steel to provide maximal internal respectively and was approximately 12 m2 in aperture area. It contains 6.5 cm deep randomly packed TiO2 coated Rashig ring and perforated bottles to direct the flow and maximize contact time.

The true quantum efficiency of the reactor ranges from 2.4-2.8 % for dichloroacetic acid mineralization. Rapid removal (96 %) of MLR was observed once exposed to sunlight with in 4 hours little difference was detected between the removal rates (70%) under cloudy or sunny conditions. The virtual independence of degradation rates for low to moderate levels of cloud cover is due to the ability of the non-concentrating reactor to harness diffuse light and minimize reflection losses.

Feitz et al (2002)

7. Batch mode plate film reactor

Diuron (810-5 mol/dm3)

Photoreactor was constructed from rectangular polymethyl methacrylate trays with a through at either end. The trays were dimensioned to accommodate a glass plate 60 cm long and 30 cm wide. The plates were positioned at an angle of 100C to the horizontal position. The plates were coated with TiO2. The flow rate was 1cm3/min

Dependence of the reaction rate on the diuron concentration (in the range of 0.8-8.010-5 mol/dm3 and the light intensity but independence on the flow rate (2.5-3.6 dm3/min) were found.

Krysova et al (1998)

51


Sl. No.

Types of Reactor



8. Novel slurry bubble column reactor

2,4-Dichloro-Phenoxyacetic acid (100 mg/L)

Continuous experiments were carried out in borosilicate glass slurry bubble column reactors of 0.1m dia. x 3.0 m length and 0.15 m dia. 3.0 m length with capacities of 20 and 54 L respectively. A metering pump was used for delivering 2,4-D solutions to the top of the column air compressor was used to sparge the air at the bottom of the column. A parabolic reflector (3.0 m height with a total surface area of 6.0 m2) was used to concentrate the solar radiation.

As the size of the reactor increase the PCO was found to decrease. An increase in the residence time yields higher photocatalytic degradation. The presence of the sieve plates with down comers increases the percentage degradation by about 10 % over that obtained in their absence. A reduction in axial mixing has a beneficial effect on the extent of photocatalytic degradation.

Kamble et al (2004)

9. Novel aerated cascade photoreactor (ACP)

Dichloro acetic acid (DCA)

It consists of a double skin sheet channeling system made of UV transparent acrylic glass. By drilling a hole at alternate ends of the channels, a meandering flow is obtained. A porous aeration tube installed at the bottom of the sheet feeds an air –stream to each channel. ACP can be installed at any angle from 0 to 900 in order to optimize the hydrodynamic state and ensure optimal irradiation of the sunlight.

For the DCA degradation using TiO2 as the catalyst, the initial photonic efficiency was 14.5 % and 23.6 % and the initial degradation rate 3.8 and 1.8 g TOC/m2/h under artificial light and sunlight respectively. When treating the textile wastewater, the removal rates of TOC and COD are almost identical with about 42 % and 29 % at an average residence time of 2 hours under artificial light and sunlight respectively.

Xi et al (2001)

52


Sl. No.

Types of Reactor

Pollutant/ Wastewaters Description of Reactor Important findings Reference

10. Parabolic Trough reactor (PTR)

Phenolic wastewater

The PTR consisted of 6 HELIOMAN-modules connected in row, its concentration ratio was 6, and the effective collecting mirror aperture area was 32 m 2. Each module was equipped with two stepper motors. Flow rates between 500 and 3000 L/h was maintained and experiments were performed in single pass mode. The total volume of the pipe reactor was 838 L of which 484 L correspond to the borosilicate glass absorber tubes.

Only 30 % of the relative TOC was degraded after 250 min. of treatment. The major engineering problem is to run the PTR maintaining a sufficient concentration of molecular oxygen throughout the reactor. Using small flow rates to achieve high residence times results in an accumulation of oxygen bubbles in the tubes thus reducing the illuminated volume and lowering the degradation rates.

Goslich et al (1997)

11. Double skin sheet reactor (DSSR)

Dichloro acetic acid (DCA)

It consisted of a flat and transparent structured box of PLEXIGLAS in which a suspension of the photocatalyst circulates driven by a pump. This reactor employs both the direct and diffuse portion of the solar radiation. It consisted of a modified double-skin sheet, length = 1400 mm, height = 980 mm, 30 channels 28.5x12 mm, total inner volume = 144 L as the photoreactor and a reservoir (30 L), connected by PVC-tubes. All experiments were operated in a recirculation mode. The volume flow was 11.8 L/min, i.e. the residence time in the reactor was 73.2 S. The flow rate being 0.57 m/s resulted in a Reynolds number of about 9000, therefore the flow conditions are turbulent. The reactor was adjusted with an inclination angle of 450C. The reactor was irradiated by the sun.

Within 60 min. more than 80 % of the initially present DCA have been degraded. Photonic efficiencies up to about 12.5 % could be achieved. Purging the suspension with molecular oxygen results in an increase in photonic efficiencies.


53


Sl. No.

Types of Reactor

Pollutant/ Wastewaters Description of Reactor Important findings Reference

12. Thin-Film-Fixed –Bed Reactor (TFFBR)

Phenolic wastewater

It consisted of a TiO2 coated glass plate with an effective surface area of 0.69m2. The plate was fixed by a metal frame and the reactor was aligned to facing south, sloping angle was always adjusted at 200 against the horizon. The flow rate was adjusted between 1 and 6 L/h, the total volume of the container was 5L. The thickness of the liquid film was between 80 and 140 m, resulting in a total illuminated reactor volume of 60-100 cm3

More than 70 % of relative initial TOC was degraded after 250 min. of treatment. In this reactor there is always a sufficient amount of oxygen present, since it is permanently present in the reactor due to the large surface area of the liquid in contact with the surrounding air.


13. Shallow pond Reactor

4-chlorophenol (8 x10 -5 M)

The experimental facility consisted of three adjacent shallow pond reactors open to atmosphere was fabricated. All three ponds rectors were initially have the same aperture 106.7 x 53.3 cm while the respective depths of the ponds were 5.1, 10.2 and 15.3 cm. The reactors were fabricated from plywood and lined with polyethylene sheets. Each pond was equipped with a mixing facility which consists of a circulating pump and a submerged spray bar. The circulating rates were 6, 10 and 18 L/min respectively.

It was found that 4CP was successfully oxidized with the photocatalyst, TiO2, suspended in slurry or adhered to a fiber glass mesh. The reactor, perform better with the slurry. It has also been found that the first-order rate constant for oxidation of 4CP increases with decreasing initial UV intensity, catalyst loading, and initial solute concentration, the oxidation rate of 4CP is invariant provided the aperture to volume ratio is fixed. It has been determined that the 4CP solution contains sufficient dissolved oxygen to support the PCO.

Wyness et al (1994 b)

54


Sl. No.

Types of Reactor



14. Flat-Plate trickle reactors

4-chlorophenol (1 10-4M)

It consisted of a rectangular stainless steel flat backing plate with dimensions 243.8 111.8 1.3 cm. A spray bar, which serves to evenly distribute the solution, is located near the top of the reactor. The spray bar has 39 evenly spaced 2.4 mm holes drilled along its length. Two 12.7 mm drain holes have been located at the bottom of the reactor.

The reaction rate constant for the slurry mode is typically two to five times greater than that for the fixed catalyst mesh tested at similar condition. In addition, the reaction rate constant appears to vary linearly with the UV isolation, and it shows no dependence on liquid film thickness in the slurry mode, but appears to vary linearly with the inverse of film thickness in the fixed catalyst mode.

Wyness et al (1994 a)

15. Rotating disk reactor

Phenol (22 mg/dm3)

The reactor consisted of three major components, a reactor vessel, a rotating disk loaded with TiO2 catalyst, and UV sources. It also included a control system for angular velocity and a controller for the UV radiation. The body of the reactor was a semicircular vessel with an inner diameter of 52 cm and a gap thickness of 3.5 cm constructed of stainless steel, had a diameter of 49.5 cm and a thickness of 0.32 cm. The capacity of the reactor was 3.5L. Four 15 W low pressure mercury UV lamps was used for artificial irradiation and solar for natural irradiation.

The solar reactor can receive solar light and oxygen from the atmosphere effectively. The phenol can decomposed rapidly under solar light than UV light. It was observed that 100 % removal was observed within 60 min.

Zhang et al (2000)

55


Sl. No.

Types of Reactor



16. Compound parabolic collector (CPC)

Synthetic Municipal Wastewater (COD = 250 mg/L)

The reactor consisted of two sunlight collectors, fixed on a platform inclined at 370. Each collector was consisting of three CPC modules mounted in series. The total surface area of each collector as 3m2. One module consists of eight parallel CPC reflectors with UV transparent tubular receivers. The three modules are connected in series and the wastewater flows directly from one to the other and finally from one to the other and finally to a tank. The total volume of the wastewater used during the experiments was 35L.

By an accumulation energy of 50 KJ/L the synergetic effect of TiO2 P 25 with H2O2 and Na2S2O8 leads to a 55% and 73% reduction of the initial organic carbon respectively. The photo-Fenton process shows to be more efficient when compared to TiO2/H2O2 system. An accumulation energy of 20 KJ/L leads to 80% reduction of organic content.

Kositzi et al (2004)

17. STEP reactor

4-Chlorphenol (45mg/L), formetanate (50 mg/L) and dye (50 mg/L)

The reactor was composed of 21 stainless steel stairs height: 70mm, width: 500 mm) covered with a 1m2 Pyrex sheet to limit water evaporation. The effluent flowed over the steps before being collected in a tank, from which it was elevated with a pump to the top of the steps for recirculation. The effluent volume, collector surface and flow rate was 25 L, 1 m2 and 450 m3/h respectively.

4-CP adsorption in the dark was higher. Complete removal of 4-CP was observed with in 1 h. Oxygen transfer was more due to formation of STEP. It was found that STEP reactor was more efficient when compared to CPC, for the removal of 4-CP and formetanate.

Guillard et al (2003)

56

2.6 BIOLOGICAL WASTEWATER TREATMENT-

SEQUENCING BATCH REACTOR (SBR)

Biological treatment has been the main technology capable of

reducing the contaminant level of wastewater for many years. The overall

objectives of the biological wastewater treatment are to transform

biodegradable compounds into acceptable end products, transform or remove

nutrients, capture suspended solids, and incorporate non-settleable colloidal

solids into biological flocs. The objective of the industrial wastewater

treatment is to remove and reduce the concentration of organic and inorganic

compounds. Although some of the organics are toxic or inhibitory to

microbial growth, a preliminary chemical oxidation step may eliminate

refractory or toxic substances. The main benefit of the biological wastewater

treatment is its relatively low operating cost and handling huge masses of

compounds (Ha et al 2000).

The successful design and operation require an understanding of

the type of microorganisms and organic compounds, the environmental

factors that affect the performance, and the types of reactors involved. The

successful operation and removal of dissolved compounds in wastewater are

done by a variety of microorganisms, principally bacteria. Microorganisms

oxidize the dissolved and particulate carbonaceous organics into simple

products and extra biomass. Among the environmental factors affecting the

treatment process, temperature, and pH have important effects on the

selection, survival, and the growth of microorganisms. The optimal growth of

a specific microorganism occurs in a fairly narrow range of temperature that

differs from one group of bacteria to the other. Most bacteria cannot tolerate

pH levels above 9.5 or below 4.0. Generally, the optimum pH for the growth

and survival of the bacteria lies between 6.5 and 7.5 (Tabrizi and Mehrvar

2004).

57

The activated sludge process has been traditionally applied to treat

industrial wastewater, but nature of such discharges often operational

problems in continuous flow systems. This is the case of wastewater

containing toxic compounds generated by several industries. In such

wastewaters, the mass of toxic pollutants could vary in time and space, thus,

to efficiently biodegrade these pollutants the treatment plants must be

designed with excess capacity. The main problem is that continuous reactors

are designed to work under steady-state conditions but, in reality, industrial

wastewater present great variability, despite equalizer tanks, giving transitory

conditions. Recently, innovative strategies like the discontinuous processes

(controlled unsteady-state processes) have been explored in order to increase

the degradation efficiencies of inhibitory wastewaters. The term sequencing

batch reactor (SBR) is used as a synonym of the wastewater treatment

technology, where the volume of the reactor tank is variable in time.

2.6.1 Sequencing Batch Reactor

The Sequencing Batch Reactor (SBR) is a fill and draw activated

sludge system for wastewater treatment. In this system, wastewater is added

to a single batch reactor, treated to remove undesirable components and then

discharged. Equalization, aeration and clarification can all be achieved using a

single batch reactor (Ketchum 1997). The sequence of operations carried out

for effective treatment involves five phases viz. Fill, React, Settle, Draw and

Idle as presented in Figure 2.4. A brief discussion of the five phases of the

SBR treatment is discussed herein under.

58

Figure 2.4 Typical operating sequences for sequencing batch reactor

2.6.1.1 FILL phase

The first phase in the sequence is filling of the reactor. The influent

to the tank may be either raw wastewater or primary effluent. Filling could be

achieved by pumping or through gravity. When a number of tanks are

operated, the wastewater is added by gravity, devices like an operated weir or

an automated valve are operated to divert the flow within the tanks. The tanks

are generally an earthen or oxidation ditch, a rectangular basin or any other

59

concrete or metal type structure. In the Fill phase, the wastewater is added to

the biomass, which is present in the tank from previous cycle. The liquid level

in tank increases from the set level to the maximum of 100 %. The set volume

may be as little as 25 % or as great as 70 %. The Fill phase can be designed to

terminate before the maximum level is reached, by either limiting the FILL

time to a predetermined maximum or eliminating IDLE phase.

Aeration is carried out during this phase by various diffusing

systems such as surface aerator and diffusers. A phase in Fill period, where no

mixing or aeration is carried out is called Static Fill. This phase utilises

minimum energy input and high substrate concentration at the end of the

FILL. Similarly, Mixed Fill imparts mixing of wastewater without aeration.

Level sensing devices, timers, or online probes can switch the aerators and/or

mixers on and off as desired. Fill is stopped, when the tank is full or diverted

into another tank in the SBR treatment scheme (Ketchum 1997).

2.6.1.2 REACT phase

Reactions are initiated during Fill, and are continued during this

phase. The react period is the time during which the tank receives no flow.

Alternating conditions of low dissolved oxygen concentration and high

dissolved oxygen concentration is adopted as per the requirement of

wastewater treatment. The liquid level remains at maximum throughout

React. Wasting of sludge could be carried out as means of controlling sludge

age. The duration of this phase can vary from zero to more than 50 % of cycle

time depending on the level of treatment achieved. React phase completes the

reactions initiated during the Fill phase (Ketchum 1997). Table 2.5 illustrated

bases of design for common operating policies to meet selected treatment

objectives.

60

Table 2.5 Common-operating strategies to meet the treatment objectives

Sl.No. Treatment Objective Fill policies React Policies

1 Organic carbon and suspended solids reduction, minimum energy consumption or sludge production

Static, mixed, then aerated

Aerated

2 Organic carbon and suspended solids reduction and nitrification


Aerated

3 Organic carbon and suspended solids reduction and denitrification


Aerated, followed by mixed then aerated

4 Organic carbon and suspended solids reduction and biological phosphate reduction

Mixed (short period), then aerated

Aerated

5 Industrial organic wastewater, toxic at high concentration

Mixed (short period), then aerated

Aerated (long period)

2.6.1.3 SETTLE phase

The reaction completed in the React phase is allowed to settle

down, wherein the solid liquid separation takes place under quiescent

conditions. Generally no inflow or outflow from the tank is practised. The

major advantage in the SBR is clarification process which is carried out in the

same tank. The biomass is retained in the tank until some fraction is decided

to be wasted. The time in Settle typically ranges between 0.5 and 1.5 h.

Sludge wasted at the end of Settle periods is harmful, as sludge may rise to

the surface (Ketchum 1997).

61

2.6.1.4 DRAW phase

This phase clearly means removal of clarified effluent. The various

mechanism for withdrawal of effluent are, i) a pipe fixed at some

predetermined level with the flow regulated with an automatic valve, ii) a

pump, depending upon the hydraulic grade line of the system, iii) an

adjustable or floating weir at or just beneath the liquid surface can be used.

The period of draw can range from 5 to more than 30 % of the total cycle time

(Ketchum 1997).

2.6.1.5 IDLE phase

The phase between Draw and Fill is termed Idle. Sludge can be

wasted effectively during Idle. The frequency of sludge wasting can be

designed to range between each cycle to once in every 2 to 3 months.

Aeration or mixing can be provided, after sludge wasting. Idle can be avoided

depending upon treatment policy by filling the tank, as soon as the tank in

Draw reaches the minimum liquid level (Ketchum 1997).

2.6.2 Design Criteria of SBR

Design of SBR systems to treat industrial wastewaters always

requires the use of treatability studies to determine appropriate SBR operating

times, including total aeration times and rates, and the need for chemical

additions for nutrients or pH control. Table 2.6 shows the design key

parameters for a wastewater system (US EPA 1999).

62

Table 2.6 Key design parameters for a conventional load

Sl. No.

Parameter Municipal wastewater

Industrial wastewater

1 Food to Mass (F:M) 0.15 – 0.4 / day 0.15 – 0.6 /day

2 Treatment Cycle Duration 4.0 hours 4.0 – 24 hours

3 Typically Low Water Level Mixed Liquor suspended Solids

2,000–2,500 mg/L

2,000-4,000 mg/L

4 Hydraulic Retention Time 6 – 14 hours Varies

2.6.2 Factors Influencing the Performance of SBR

Several physical parameters have been shown to affect the

performance of SBR. The main parameters include temperature, long idle

periods and oxygen. Many works have been done to investigate the effects of

such parameters on the SBR kinetics.

2.6.3.1 Effect of temperature

Temperature is one parameter, which affects the performance of

SBR. The performance of SBR is different under identical loading conditions.

Higher removal efficiency could be obtained at temperatures ranging from

20 – 35˚C. Large number of filamentous organisms was produced at high

temperatures. However higher removal efficiency could also be obtained at

lower temperatures if wastewater is diluted. Biological treatment of

wastewater at higher temperature has the advantage of providing effective

treatment with minimum energy and operating cost. Chemical compound

analysis of feed and various treated effluents showed a better removal of long

chain fatty acids at thermophilic conditions (60˚C) in comparison to

mesophilic conditions (35˚C) (Ndon et al 1997 and Tripathi et al 1999).

63

2.6.3.2 Effect of long idle periods

Sludge storage can be used as an effective control to adjust plant

capacity to large influent variations. The storage of sludge can be done in an

external tank or in the mainstream system. Mainstream storage is possible in a

treatment plant with parallel treatment trails, where individual trains can be

individually switched on and off. The SBR technology is well suited for this

temporary sludge storage because reactors can easily switched off

individually and operated in an idle mode. A number of researchers have

investigated the effect of storing bacteria without external substrate.

Morgenroth et al (2000) conducted batch experiments to find the effect

of long term idle period on the performance of SBR. SBRs were operated for

idle times ranging from 6 to 20 days and was found that after 6 weeks of

storage and without aeration the sludge still had above 55 % of the initial

volumetric ammonia oxidation rate. No adverse effects on sludge settleability

were observed with a SVI below 140 mL/g in the lab scale SBRs. In the

laboratory scale activated sludge sequencing batch reactor idle periods had no

effect on nitrogen removal. For longer idle periods ammonia oxidation was

not significantly affected but reactor operation had to be adjusted for nitrite

accumulation, time for denitrification was to be increased by supplying an

external carbon source. Also with the increase in time of idle more quantity of

ammonia was oxidised to nitrate.

2.6.3.3 Effect of oxygen supply

Oxygen supply methods also affected the performance of SBR.

Dangcong et al (2000) studied high ammonium nitrification studies in SBRs

by two oxygen supply methods, viz. Controlled and uncontrolled oxygen

supply method. In the controlled oxygen supply method dissolved oxygen

64

concentration ranged between 2-3 mg/L. However, in uncontrolled method,

the concentration of DO changes with the oxygen utilisation rate. The

experimental results showed that oxygen supply methods have significant

effects on SBR performance. High ammonia nitrogen oxidation took place

simultaneously with the uncontrolled oxygen supply method. Ammonium

oxidation proceeded smoothly during the low DO period of the cycle,

however nitrate oxidation was completely inhibited at low DO but recovered

during the high DO period of the same cycle. Ammonium oxidation

developed an ability to sustain the fluctuations of DO, but nitrate oxidisers did

not.

2.6.4 Devices for Aeration, Mixing and Decanting

2.6.4.1 Aeration

Aeration is the process by which oxygen is introduced into the

system for the oxidation of organic and inorganic wastewater compounds. In

addition to the transfer of oxygen into the water, aeration usually provides

sufficient mixing and removal of reaction products (gases) from the water. A

number of devices/methods are available to aerate. Some of the devices are

surface aerators, diffusers, and jet aeration. In selection of aeration device,

two aspects are of importance, energy efficiency and reliability. An energy

efficient aeration device should be selected, as aeration is among the most

energy intensive operations in wastewater treatment systems, consuming

50-90 % of the total energy cost (Edgertor et al 2000).

2.6.4.2 Mixing

Mixing of wastewater components (substrate) to the

microorganism’s needs is to be ensured by mixing devices. Some of the

65

mixing devices used are jet mixer, surface rotor and a submerged propeller.

Mixer devices should be selected to achieve good mixing with limited

introduction of oxygen due to turbulence at the surface (Kazmi and Furumai

2000).

2.6.4.3 Decanting

A variety of decanters are used to remove the treated supernatant

from the reactor. Three categories are distinguished. The decanters remove

the supernatant from the surface of the reactor by a motor to rotate the trough

with the overflow weir, the speed of rotation to lower the weir is used to

control the rate of draw. In another method decanter floats on the surface of

water and withdraw rate is determined by the hydrostatic pressure at the inlet

of the decanter. In some instances, decanter is installed at a fixed depth in the

SBR. When designing a decanter, precautions to account for foaming sludge

must be included (Morgentath et al 2000).

2.6.5 Comparison of SBR with Conventional Continuous Flow

Processes

SBR and the conventional continuous flow system (CFS) differ in a

way that former provides in time, what the CFS provides in space. SBR can

meet even more stringent effluent criteria that are being promulgated, since

they are operated in batches and flexibility to treat on merit of the wastewater.

In fact a properly designated semi batch reactor should achieve a higher

effluent quality than CFS. In contrast to continuous flow system, both

biological reactions and biomass separation take place in the same tank. This

is a single tank batch system, sequencing on individual cycle and thus

provides both low capital and operating costs. Conventional waste

66

management technologies commonly adopted in tropical climates are not only

expensive but also warrant exacting operation and maintenance requirements.

In addition to this one more important parameter associated with

SBR, and not found in CFS is flexibility of operation. The time of all the

phases (fill, react, aeration etc.) can be easily changed which facilities the

SBR to tolerate variations on organic and hydraulic loading and to sustain the

shock loads. After the react phase, the microbial sludge is allowed to settle in

the same tank which has comparatively a larger volume than that of secondary

clarifier used for conventional continuous flow activated sludge plant. SBR

has the characteristics ability to incorporate aerobic and anoxic phases in a

single reactor, if desired.

There are many continuous flow systems in which the air supply

has been cycled on and off in an attempt to provide for both nitrification and

denitrification without the addition of an external supply of substrate (for

example methanol). The balance between air supply and external substrate

concentration in a continuous flow system with or without recycle is quite a

sophisticated operational problem, which is believed to be unsuccessful in

developing countries. A batch reactor on the other hand can be easily

operated. In the SBR process, during each cycle dynamic conditions are

prevailing. Control and periodic repetition of the short term unsteady state

allow for, in the long run, enhancement of certain effects such as enzymatic

activity, accumulation of metabolic products and selection and enrichment of

specific groups of microorganisms. Similar effects are achieved in continuous

flow activated sludge system with a plug flow or cascade flow bioreactor.

The SBR has the advantage of a more flexible operation as stated

earlier with true plug flow characteristics. Since the treatment process

proceeds over time in one reactor, the process control can be conventionally

executed by means of auto-analyzers (e.g., for dissolved oxygen, pH, nitrate)

67

and computers. On the other hand, continuous flow systems are more flexible

with respect to recycle of sludge and wastewater. Recycle flows can be

directed to any reactor in the system where as the comparable operation is not

possible in the SBR. SBR treatment profile functions similarly as in a

continuous flow stirred tank reactors scheme, where the concentration of

substrate reduces in similar pattern. In a similar manner for anaerobic

treatment of wastes, anaerobic sequencing batch reactor (AnSBR) is utilized.

The commonly used mode for anaerobic digestion is the Upflow Anaerobic

Sludge Blanket reactor which is continuously fed in an up-flow mode and the

influent and effluent must be fed and withdrawn uniformly across the

horizontal plane of the reactor. In contrast, the AnSBR is intermittently fed.

An elaborate wastewater feed and gas/effluent collection system is not

required for the AnSBR, since AnSBR is batch-fed and drawn, there is no

possibility of short circuiting of waste through the reactor and no

re-circulation of effluent is needed. The comparison between the sequencing

batch reactors and the conventional continuous flow system is given in

Table 2.7.

Extensive lab and pilot scale studies have shown that various types

of wastewater could be treated using sequencing batch reactor. It also

provides a viable alternative to continuous flow systems in nutrient removal

as well as in carbon and suspended solids removal. It offers a great deal of

operational flexibility, as it allows for easy adjustment of aerobic, anoxic and

anaerobic periods through temporal control of aeration and filling.

Application of automatic control to SBR for cost reduction and process

optimization are some of the main causes of renewed interest in the SBR

technology. Because of the flexibility associated with working in time, the

SBR can be operated as either a labour-intensive, low-energy, low- sludge

yield system or a minimal labour, high-sludge yield system for the same

physical plant. The actual operation policy can be adjusted in accordance with

68

economic condition by modifying the settings of control mechanism. Labour,

energy and sludge yield can also be traded off with initial capital costs

(Pratibha et al 2004).

Table 2.7 Comparison between SBR and continuous flow reactors

Sl. No. Parameter Batch Process Continuous Process

1 Concept Time sequence Spatial sequence

2 Inflow to reactor Periodic Continuous

3 Discharge from reactor

sequenced Continuous

4 Aeration sequenced Continuous

5

Mixed liquor

Always in reactor

No recycle

Recycle through reactor and clarifier

6 Clarification Quiescent hydraulics

Hydraulic motion

7 Flow pattern Perfect plug Complete mix or approaching plug

8 Equalization Inherent None

9 Flexibility Considerable Limited

10 Hydraulic sizing Variable Uniform

11 Capital cost Low Medium

12 O & M cost Low Low

13 Possible automation Yes No

14 Physical structure Compact Larger than SBR

15 Sludge yield Low Moderate

16 Effluent quality Good Variable

69

2.7 OVERVIEW OF WORK DONE ON BIOLOGICAL

TREATMENT

Biological treatment of phenolic wastewater is generally favoured

over other processes in terms of lower cost as well as the possibility of

complete mineralization. Some of the illustrative work done in the recent

years in the area of biological treatment has been furnished in Table 2.8. Chan

et al (2006) studied degradation of phenol containing wastewater by SBR. He

observed the inhibitory effect of phenol on bioactivity when the concentration

of phenol increased from 100 to 1000 mg/L. Sahinkaya et al (2006)

investigated effect of biogenic substrate concentration on the performance of

SBR treating 220 mg/L 4-CP and 110 mg/L 2, 4-DCP mixture. It was

observed that decreasing peptone concentration associated with decreasing

biomass concentration led to the observation of lower degradation rates,

which caused accumulation of chlorophenols within the reactor.

Accumulation of chlorophenols further decreased the removal rate due to self

inhibitory effect of chlorophenols on their own degradation and strong

competitive inhibition of 2, 4 DCP on 4-CP degradation. Ha et al (2000)

studied COD removal of phenolic wastewater by biological sequencing batch

reactor in the presence of 2, 4-DCP. It was observed that increase of influent

concentration by 10 % resulted in increased COD concentration from 37.7,

19.6, and 14.5 mg/L to 51.3, 28.1 and 18.1 mg/L.

Annadurai et al (2002) studied degradation of phenol using mixed

liquors of Pseudomonas putida and activated sludge. It was observed that for

an initial concentration of 500 mg/L, only 22 % of phenol is degraded using

activated sludge for 48 h. Kumar et al (2005) studied degradation of phenol

and catechol using Pseudomonas putida MTCC 1194. The well acclimatized

culture of Pseudomonas putida degraded the initial phenol concentration of

1000 mg/L and catechol concentration of 500 mg/L completely in 162 and

70

94 h respectively. Maeda et al (2005) observed that o-cresol concentrations

greater than 30 mg/L were toxic to the microbial activites. Kargi and Konya

(2006) studied the degradation of para-chlorophenol in an activated sludge

unit. He concluded that percentage removal of COD and 4-CP decreased with

increase in the feed 4-CP concentration above 925 mg/L, due to inhibition

effect. Many researchers Lourenco et al (2000), Vijayaraghavan et al (2007),

Farabegoli et al (2004), Mangat et al (1999), Yoong and Lant (2001), Erkan

et al (2006), Tripathi and Allen (1999) have observed the same effect for

degradation of reactive textile dye, palm oil mill effluent, 2,

4- dichlorophenoxyacetic acid, phenolic wastewater, tannery, 4-CP, pulp mill

effluent. Hence, there is a limitation associated with the biological treatment

due to the toxicity exerted by phenol itself at higher concentration. SBR

suffered by several constraints when used for toxic wastewater degradation,

inhibition of the microorganisms, problems with shock loads of toxic

compounds, deacclimation of the microorganisms and low efficiencies

regarding the removal of toxic compounds.

Shock loads appear when the toxic concentration in the influent

greatly increases as a consequence of changes in the manufacturing process,

for example, during the cleaning of the production units. A toxic shock

produces an increase in reaction time, diminishes the efficiency and may

posion the biomass (Watanabe et al 1996). Research efforts have been made

to overcome the difficulties associated with the treatment of an inhibitory

compound like phenol. The use of advanced oxidation processes in

conjunction with the biological oxidation has been a recent innovation in the

treatment strategies for wastewater treatment.

71

Table 2.8 Typical findings observed in the representative works related to biological treatment

Sl. No.

Compound/wastewater Experimental condition and Type of equipment Important findings of the work Reference

1. Bleached Kraft mill wastewater (COD=1400 mg/L)

Two activated sludge bioreactors were constructed out of 6 mm thick plexiglass. The height of each reactor was 25.5 cm and inner diameter was 19 cm. The inner volume 10 L contained 5 L of activated sludge. Aeration was supplied by passing compressed air. The MLSS was maintained in the range of 1500-2000 mg/L

At steady state (HRT 10-12 h, SRT 12-15 d), the percentage removal of BOD, COD, and toxicity averaged 87, 32, 97 respectively. Varying HRT between 12 and 4 h and SRT between 5 and 15 d indicated that HRT had more of an effect on treatment performance than SRT. Longer HRT (12 h) led to improved BOD (87 %), COD (50 %), toxicity (90%) removal, while longer SRTs were not shown to significantly affect the performance.

Barr et al (1995)

2. Piggery waste water (TOC = 5860 mg/L)

The anaerobic reactors had an active volume of 1.5 L. The temperature was kept constant at 350 C. Two aerobic reactors were used in this study. The first one had an active volume of 1.5 L. When the organic carbon load of the system was double, a 4 L reactor capacity was used. Aeration was provided by compressors to plastic tube placed at the bottom of the reactors.

Combined anaerobic-aerobic system was investigated using two Lab-scale sequencing batch reactors. The cycle length was 24 h. Three recycle-to-influent ratios from 1 to 3 were tested. Average performances of the overall process, in the different conditions tested, were a TOC removal of 81 to 91% and 85 to 91 % of TKN.

Bernet et al (1999)

3. Oil shale ash leachate (COD = 2000-4600 mg/L)

The anaerobic reactors were glass upflow anaerobic sludge blanket reactors (height 24 cm, diameter 3.2 cm) with total liquid volume of 0.2 L. The aerobic reactors were plastics activated sludge reactors, each with n aeration unit (liquid volume 0.2 L) and a separate settling unit (liquid volume 0.2 L). Aerobic conditions (>2 mg/L) and mixing in the aerobic reactors were assured through continuous aeration with submerged diffusers in each vessel.

In the sequential anaerobic – aerobic processes, the COD removals were most of the time slightly lower at 100C (67 %) than at 200C (78 %) while the BOD for 7 days removals was 97-99 % at both temperatures. In the single aerobic process at 200 C, the COD removal was 65 % while at 70 C the COD removal was 54 %. After the feed was changed from leachate to phenol in the 70 C aerobic reactors, the COD removal stabilized to about 95 %. In all the leachate treatment processes studies, the total phenols removals were on average 78-86 %. The anaerobic stages removed total phenols minimally

Kettunen et al (1996)

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Sl. No.


4. Paper mill process water (COD=2113 mg/L)

Two 4 L stainless steel activated sludge reactors were used. Both reactors were divided in five compartments, each compartment containing 800 mL of mixed liquor. Each compartment was aerated with pressurized air, sparged through an aeration stone plug flow was maintained. COD: N: P = 100:2.6:0.4

Two lab-scale plug flow ASP reactors were run in parallel for 6 months, a thermophilic reactor at 550 C and a reference reactor at 300 C. Both reactors were operated simultaneously at 20,15 and 10 days SRT. COD removal percentages were 58 5 % at 300C and 48 10 % at 550 C. The effect of the SRT on the total COD removal was negligible. At 300 C, COD removal percentages were 65 %, 75 % and 86 % at 20, 15, and 10 days SRT respectively. At 550C, these percentages were 48 %, 40 % and 70 % respectively. The mesophilic activated sludge showed a higher removal of total COD as compared to the thermophilic reactor.

Vogelaar et al (2002)

5. Pulping whitewater (COD=2500 mg/L)

A lab-scale Plexiglas reactor with a total liquid volume of 8.55 L was used in the study. The reactor was filled with 58 % Kaldnes carrier elements, occupying 11 % of the reactor’s liquid volume. The reactor was kept in a kept in a temperature-controlled water batch at 550C. Mixing and aeration were provided by pressurized air through ceramic aerators in the bottom of the reactor.

The continuously operated Lab-scale Kaldnes moving bed biofilm reactor (MBBR) was used for aerobic treatment of pulping white water. Inoculation with mesophilic-activated sludge gave 60-65 % COD removal from the first day onwards. During 107 days of experiment the 60-65 %. SCOD removals were achieved at organic loading rates of 2.5-3.5 Kg COD/m3/d, the highest loading rates applied during the run and HRT of 13-22 h.

Jahren et al (2002)

6. Sugar beet wastewater (COD=2500 mg/L)

Lab-scale mesophilic (20-350C) and thermophilic (550C) ASP was made of PVC and consisting of aeration tank of volume 1.5 L and settling units (0.4 L). Aquarium aerators were used for aeration. SRT was 116 d in the mesophilic ASP and 53 d in the thermophilic ASP. The MLSS and MLVSS were 5.6 g/L and 3.7 g/L respectively.

In the ASPs, the HRT was 12 h in both processes, corresponding to a volumetric loading rate of 3.2 1.0 kg COD/m3/d. The Mesophilic ASP gave 79 18 % and the thermophilic ASP gave 50 6 % of COD removals. Both ASPs gave 90 % COD removals at 24 h HRT.

Suvilampi et al (2005)

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7. Phenol (500 mg/L)

Experiments were performed in a column type SBR @ 250C. The reactor had a working volume of 20L, with an internal diameter of 50.0 mm. Air at a flow rate of 3.5 L/min was introduced by a fine bubble aerator in the bottom of the column. The reactor was operated sequentially in 4 h cycles.

The reactor was started at a loading rate of 1.5 kg phenol/m3/d with phenol-enriched activated sludge as inoculums. The phenol loading was then adjusted stepwise to a final value of 2.5 kg phenol/m3/d. At this high loading, phenol was completely degraded and high biomass concentration was maintained in the rector. High specific phenol degradation rates exceeding 1 g phenol g/VSS/d were sustained up to phenol concentration of 500 mg/L and significant rates continued to be achieved up to a phenol concentration of 1,900 mg/L.

Tay et al (2004)

8. Paper mill wastewater (COD=1000 g/m3)

Four-bench scale SBRS were fed with pre-settled wastewater from a paper mill. The volume of each reactor was 6.5 L. All reactors were operated with the total cycle time of 8 hours, i.e. three cycles daily. In addition a fourth reactor was operated with total cycle time of 24 hrs. The COD: N: P was 100:3:1. MLSS was maintained in the range of 2600-5000 mg/L.

The influence of process conditions, like duration of filling phase (Fill) or duration of reaction phase (REACT) could be demonstrated. The highest COD removal of 92 % and the best sludge settling properties for the paper mill wastewater were obtained at a sludge age of 20 days with REACT period of 12 hours and duration of FILL phase of 0.5 hours.

Franta et al (1996)

9. Phenolic wastewater (1300 mg/L)

The bench-scale reactor comprised a Perspex cylinder with a detachable base-plate secured to one end with screws and an O-ring seal. The top of the cylinder was enclosed with a removable cover in two halves. Aeration was provided at the base of the cylinder through two air-stones at opposite ends. A magnetic stir-bar ensured complete mixing in the reactor. MLSS was maintained in the range of 3500-3900 mg/L.

The investigation demonstrated the capability of a bench-scale sequencing batch reactor (SBR) to biodegrade an inhibitory substrate at a high loading rate of 3.12 kg phenol/ m3/d (2.1 g COD /g MLVSS/ d) with a COD removal efficiency of 97 % at a SRT of 4 days and a HRT of 10 hours was achieved. The SBR was operated at 4 hours cycle, including 3 hours react phase.

Yoong and Lant (2001)

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Sl. No.

Compound/wastewater

Experimental condition and Type of equipment

Important findings of the work Reference

11. 2,4-Dichloro-Phenoxyacetic acid (100 mg/L)

Four identical lab-scale SBRs were used. Each rector was made of 6 mm thick Plexiglas cylinders with an internal diameter of 100 mm. The operating liquid volume was 2.0 L. Aeration was supplied through submerged diffuses placed at the bottom of the rectors. Temperature was kept constant at 220 C. The SRT was kept constant at 20 days HRT was varied in the range of 12-48 hours.

A bench scale study using SBR was conducted to investigate the effects of HRT, the presence or absence of supplemented substrate and variation in feed concentration on the degradation of 2,4-D was studied. A long acclimation period (about 4 months) was observed before 2,4-D biodegradation was established. At steady-state operation, all reactors achieved complete removal (>99 %) of 2,4-D and the corresponding supplemental substrate, regardless of the HRT applied, ranging from 12 to 48 h. The 2, 4-D specific removal rates were affected by the type of substrate used (Phenol or dextrose) being significantly lower (30 to 50 %) in the case of dextrose.

Magnat and Elefsiniotis (1998)

12. Tannery soak liquor (COD =1500 – 3600 mg/L)

The lab-scale SBR had a volume of 10 L. An air compressor delivering airflow of 1.2 L/min supplied aeration. Each cycle lasted for 24 h, the reaction took place in 22 h, the settling in 1 h 30 min and the withdrawal and filling of the treated effluent and influent in 30 min.

Once the acclimation of the microorganisms was achieved. Optimum removal efficiencies of 95 %, 93 %, 96 % and 92 % on COD,PO4

3-, TKN and SS, respectively, could be reached with 5 days hydraulic retention time (HRT), an organic loading rate (OLR) of 0.6 kg COD m3/d and 34 g NaCl/L. The organisms responsible for nitrogen removal appeared to be the most sensitive to the modifications of these parameters.

Lefebvre et al (2005)

13. Brewery wastewater (COD= 212 mg/L)

Experiments were carried out in a cylindrical column reactor (110 cm filling height, 10 cm dia.) with a working volume of 8.6 L. Aeration was provided by means of air bubble diffusers at a volumetric flow rate of 500 L/h. Temperature was controlled at 25 ± 2 _C. The reactor was operated in SBR mode with total cycle duration of 6 h.

Aerobic granular sludge was cultivated in a sequencing batch reactor fed with brewery wastewater. After nine-week operation, stable granules with sizes of 2–7 mm were obtained. After granulation, high and stable removal efficiency of 88.7 % COD was achieved at the volumetric exchange ratio of 50% and cycle duration of 6 h. The average total COD and soluble COD of the effluent were 212 and 134 mg/L, respectively, and the average effluent ammonium concentration was less than 14.4 mg/L. Nitrogen were removed due to nitrification and simultaneous denitrification in the inner core of granules.

Wang et al (2007)

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Compound/wastewater


14. Tannery wastewater (COD = 4800 mg/L)

A bench-scale SBR made of plexiglass and of working volume 8 L was used for the study. Aeration was provided from a diffuser at the base of the reactor.

Measurement of oxygen uptake rates (OUR) and corresponding COD uptake rates showed that a 12 h operating cycle was optimum for tannery wastewater. At a 12 h SBR cycle with a loading rate of 1.9– 2.1kg/m3/d, removal of 80–82 % COD, 78–80 % TKN and 83–99 % NH3-N were achieved. These removal efficiencies were much higher than the conventional aerobic systems.

Ganesh et al (2006)

15. 4-chlorophenol (200 mg/L)

A fermenter with 5 L working volume was used as the sequencing batch reactor (SBR). Aeration was provided by using an air pump. Agitation speed was varied from 25 to 300 rpm depending on the operation phase.

Percent nutrient removals increased with increasing sludge age and decreasing 4-CP concentrations. Low nutrient removals were obtained at high initial 4-CP concentrations especially at low sludge ages. However, high sludge ages partially overcome the adverse effects of 4-CP and resulted in high nutrient removals. COD, NH4-N, PO4-P and 4-CP removals were 76 %, 72 %, 26 % and 34 % at a sludge age of 25 days and initial 4-CP concentration of 200 mg /L. Sludge volume index (SVI) also decreased with increasing sludge age and decreasing 4-CP concentrations.

Kargi and Konga (2006)

16. Paper factory effluent (COD = 1100 mg/L)

Continuous treatment of the phenolic effluent was performed with a packed bed reactor in a glass column (30 cm length and 6 cm diameter). Immobilized cells in beads were packed to a height of 25 cm in the glass column. The reactor was fed with effluent at varying flow rates of 2.5, 5.0, 7.5 and 10 mL/h.

Treatment of the paper factory effluent was done with free and immobilized cells of a phenol degrading Alcaligenes sp. d2. The free cells could bring a maximum of 99 % reduction in phenol and 40 % reduction in COD after 32 and 20 h of treatment, respectively. In the case of immobilized cells, a maximum of 99 % phenol reduction and 70 % COD reduction was attained after 20 h of treatment under batch process. In the continuous mode of operation using packed bed reactor, the strain was able to give 99 % phenol removal and 92 % COD reduction in 8 h of residence time The optimum flow rate was 2.5 mL/h and the half life period was 76 h.

Nair et al (2007)

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17. Electroplating wastewater ( COD = 5l6 mg/L)

Six 10 L reactors made from acrylic plastic were used. The size of each reactor was 18 cm diameter and 40 cm height. The total volume and working volume of the reactor were 10.0 and 7.5 L, respectively.

SBR system showed the highest COD, BOD5, TKN and cyanide removal efficiencies of 79 %, 85 %, 49 % and 97 %, respectively with 4-times diluted rice noodle wastewater (RNWW) containing 10% EPWW and 138 mg/L NH4Cl (BOD5: TN of 100:10) at SRT of 72 ± 13 days (under organic and cyanide loadings of 0.40 kg-BOD5/m3 d and 0.0023 kg-CN/m3 d, respectively).

Sirianuntapiboon et al (2007)

18. Simulated textile wastewater (COD = 4200 mg/L)

A continuously fed stainless steel anaerobic UASB and aerobic CSTR reactor were used in sequence for the experimentation. The UASB reactor had 2.5 L of effective volume with an internal diameter of 6 cm and a height of 100 cm. The CSTR reactor consisted of an aeration tank (effective volume = 9 L) and a settling compartment (effective volume = 1.32 L).

Hydraulic retention times were changed to determine the effect of HRT on removal efficiencies of colour, COD and total aromatic amine (TAA) through 186 days containing 46 days of steady-state and acclimation periods. COD and colour removal efficiencies varying between 97 % and 91 % and between 84 % and 91 % were obtained at a total HRT of 19.17 and 1.22 days in combined anaerobic/aerobic system, respectively. In the sequential aerobic stage the significant part of TAA was removed successfully while the colour removal slightly increased with TAA removal efficiencies of 70–85 % at total HRTs of 8.85 and 6.05 days, respectively. Increases in HRT provide enough time for partial mineralization of COD and intermetabolites in anaerobic and/or anaerobic/aerobic systems.

Isik et al (2008)

19. Textile wastewater (COD =1000-4000 mg/L)

The aerobic system used was a combined CSTR and FFB bioreactor. The continuous stirred tank reactor with a 700 mL working volume was used. The FFB is a 1.5 L continuous flow with three compartments packed with a rippled cylindrical polyethylene support.

This system gives high COD and colour removal efficiencies of 97.5% and 97.3 %, respectively, obtained with a total hydraulic retention time (HRT) of 4 days and total OLR of 0.29 g / L /d by the sloughing of biofilm, and the washout phenomena.

Khelifi et al (2008)

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20. Phenol ( 250 mg/L)

The pre-denitrification system consisted of an anoxic upflow sludge blanket reactor (0.8 L) and an aerobic activated sludge reactor (1.8 L). The system was provided with a liquid displacement biogas measurement device.

The effect of phenol overloads on the removal of organic matter and nitrogen compounds was studied. During the overloads from 250 to 4000 mg/L, phenol was detected in the effluent of the anoxic reactor but the system recovered fast after stopping the overloads. TOC removal remained unchanged during phenol addition (91.9 % at 0.20 kg TOC/m3/d), except for the highest overload. Phenol concentrations from 250 to 4000 mg/L were added to the feed. Phenol was completely removed despite the presence of other carbon sources in the wastewater. In spite of the presence of phenol, a TOC removal around 91.3 % was achieved at an average organic loading rate of 0.11 kg TOC/m3 d.

Eiroa et al (2008)

21. Paper mill effluent

Four identical 4 L reactors were constructed with acrylic boards. MLSS concentration was between 3000 and 5000 mg/L with a sludge age of 5-10 days.

Laboratory scale research on the effects of operating parameters, including mixed liquor suspended solid (MLSS) concentration, volumetric exchange rate (VER), aeration time, temperature and daily operation cycle on biological treatment of the pulp and paper mill effluent was studied using four 4 L sequencing batch reactors (SBR). The results showed that COD removal efficiency was up to 93.1±0.3 % and the volumetric loading reached 1.9 kg BOD m3 /day under optimal operating conditions.

Tsang et al (2007)

22. Fish market wastewater

The four reactors were made of acrylic cylinders having an outer diameter of 24 cm and a height of 34 cm. The effective working volume was 10 L, with 6 L of biomass. SRT was varied between 20 and 100 days.

The effects of COD/N ratio (3 – 6) and salt concentration (0.5 – 2 %) on organics and nitrogen removal efficiencies in three sequencing batch reactors (SBRs) with synthetic wastewater and one SBR with fish market wastewater were investigated under different operating schedules. The solids retention time (SRT, 20 – 100 days) and aeration time (4 – 10 h) was also varied to monitor the performance. For synthetic wastewater, COD removal efficiencies were consistently greater than 95 %, irrespective of changes in COD/N ratio, aeration time and salt concentrations. Increasing the salt concentrations decreased the nitrification efficiency, while high COD/N ratio’s favored better nitrogen removal (>90 %). The treatment of real saline wastewater (>3.2 %) from a fish market showed high COD (>80 %) and nitrogen (>40 %) removal efficiencies despite high loading rate and COD/N fluctuations, which is due to the acclimatization of the biomass within the SBR.

Rene et al (2008)

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23. Combined pharmaceutical and tannery wastewaters (COD = 2200 mg/L)

The 4 L fill-and-draw reactor, equipped with diffused aeration devices was operated at steady state with a sludge age of 10 days and a hydraulic retention time of 1 day.

The tannery sample was a plain-settled effluent having a total COD of around 2200 mg/L with a readily biodegradable fraction of 15 %. The same fraction was 57 % in the pharmaceutical wastewater sample having a much stronger total COD content of 4043 mg/L. Consequently, mixing of the pharmaceutical effluent with the tannery wastewater up to 38 % of the total COD in the mixture increased the readily biodegradable COD fraction but had an inhibitory effect on the biodegradation kinetics. This effect was relatively lower on growth, but quite significant on the hydrolysis of the slowly biodegradable COD decreasing the maximum hydrolysis rate from 2.0 day−1 to 1.2 day−1.

Cokgor et al (2008)

24. o-cresol (100 – 600 mg/L)

A slurry batch reactor of capacity 2 L was used as bioreactor. Aeration was provided at the base of the reactors. MLSS was maintained in the range of (2000 – 3000 mg/L)

The biodegradation kinetics of o-cresol was examined by varying initial o-cresol concentrations (30 – 600 mg/L), MLSS (1000 – 11500 mg/L) and aeration rates (0.05 – 1 L/min). It was found that about 7 h were required to biodegrade a phenol of 600 mg/L completely. The biodegradation rate was independent of the aeration rate Q ≥ 0.25 L/min. when Q ≤ 0.25 L/ min oxygen supply by the aeration was not sufficient to degrade the O-cresol.

Maeda et al (2005)

25. Phenol (100 – 1000 mg/L)

Batch reactor of dimensions 20 x 15 x 25 cm with total working volume of 5 L operated at a cycle time of 12 h .MLVSS was maintained in the range of 2000 – 4000 mg/L

It was observed that the inhibitory effect seemed to be more pronounced with the increase in the influent phenol concentration from 100 – 1000 mg/L. The k value was reduced from 290 x 10 -3 min-1 to 4.3 x 10 -3 min-1 when the phenol concentration was increased from 100 mg/L to 1000 mg/L. The phenol removal efficiency of 99 % was observed.

Chan et al (2007)

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26. Reactive textile dye (60 – 100 mg /L)

The experimental system used was composed of two 1 L reactors operating in a sequencing batch mode (SBR) in 24 hour cycle with five discrete periods fill – 50 min, react-21 h , settle -1 h , draw -55 min and idle -15 min.

When comparing with sludge retention time of 15 days the total COD removal was around 80 % with 30 % removed anaerobically. After 40 – 50 days of acclimatization the colour removal efficiency reached a maximum, stable value of 90 % from a feed by concentration of 90 mg/L. Reduction of the SRT value from 15 to 10 days reduced the biomass concentration from 2.0 to 1.2 VSS/L and lowered colour removal levels from 90 % to 30-50 % then the SRT value was increased back to 15 days the colour removal efficiency of the system was completely recovered, suggesting that with a SRT of 10 days the adequate microbial population could not be installed in the reactor.

Lourenco et al (2000)

27. Palm oil mill effluent (PoME) ( COD = 1000 – 5000 mg/L)

The activated sludge reactor had a rectangular shape with 650 mm length x 400 mm width x 350 mm liquid depth. The aeration was carried out using a diffused aeration system. The MLSS was maintained in the range of 3900 - 4200 mg/L.

The initial studies on the efficiency of the activated sludge reactor were carried out using diluted raw PoME for varying the HRT viz. 18, 24, 30 and 36 h and influent COD concentration, viz , 1000, 2000, 3000, 4000 and 5000 mg/L , respectively. The results showed that at the end of 36 h HRT, the COD removal efficiencies were found to be 83 %, 72 %, 64 %, 54 % and 42 % where at 24 h HRT, the COD removal to 57 %, 45 %, 38 %, 30 % and 27 % respectively.

Vijayaraghavan et al (2007)

28. Phenolic wastewater (COD = 480 mg/L & 560 mg/L)

A cylinder shaped reactor of acrylic plastic with 12 L maximum capacity and 175 mm in diameter and 500 mm in height was used. Air was supplied to the reactor through the air diffuses and controlled by solenoid valve. SBR cycling time was 8 hours

Treatment performance of the mixture of phenol and 2, 4 – DCP by the BAC-SBR and the SBR systems was studied by changing influent concentration and SRT. By increment of the influent COD concentration from 480 to 560 mg/L, the COD removal efficiency in BAC-SBR decreased from 92.1, 95.9 and 97.0 % to 90.8, 94.9 and 96.8 % for SRTs of 3, 5 and 8 days respectively. Meanwhile the COD removal efficiency in the SBR system changes from 85.9, 91.6 and 95.1 % to 83.6, 92.2 and 95.3 % for SRTs of 3, 5 and 8 days. When the SRT was increased from 3 to 8 days the COD removal efficiency was increased from 90.8 – 97 % in BAC-SBR and 83.6 – 95.3 % in the SBR.

Ha et al (2000)

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29. Kraft mill wastewater

Aerobic reactor of 4.5 L capacity was used. The OLR was varied from 0.4 to 1.42 g BOD/L/d and HRT was varied from 48 to 4.5 h SRT varied between 25 and 30 days

The result showed that high BOD (90 %) and COD (58 %) removal was observed when HRT varied from 16 to 6 h. Degradation of phenolic compounds was seriously affected by HRT variation, obtaining the highest removal efficiency at HRT of 48 h (33.5 %) and a minimal efficiency at HRT of 4.5 h (3.6 %). When HRT less than 6 h, the system showed destabilization and COD, BOD and SS removal decreased.

Diez et al (2002)

30. Bleached Kraft mill effluent (COD=880 mg/L)

Aerated Lagoon with an aerated (0.44 L) and a settling zone (0.22 L) was used as biological treatment. HRT was 44.5 and 45.2 h.

An aerobic lagoon was able to remove over 98 % and 80 % of the BOD and abietic acid respectively, when operated at HRT of 2 days. Under these conditions, the COD removal efficiency of 67.3 % was observed

Belmonte et al (2006)

31. Coke wastewater (COD = 3275 mg/L)

Aerobatic reactor consisted of 20 L volume made up of PVC. Oxygen was introduced through the bottom. MLVSS was maintained in the range of 1000 – 2000 mg/L

When bicarbonate was added, the maximum removal efficiencies of 71 %, 65 % and 97 % were obtained for ammonium nitrogen, COD, Phenols respectively for HRT of 54.3 h. When bicarbonate was not added this efficiency were 71%, 75% and 98% respectively. The biodegradation of phenol (504 mg/L) improves with increasing pH, achieving 96 % at pH 8 in 15 h.

Vazquez et al (2006 a)

32. Coke wastewater (COD 922-1980 mg/L) (Phenol=293 mg/L)

A Lab-scale biological plant composed of two aerobic reactors operating at 35˚C. HRT was varied from 27-98 h.

When an effluent recycling ratio of 2 is employed, average removal efficiencies of 86.2, 98.8, 97,9 and 99.3 % for COD, phenol, SCN- and NH4

+-N respectively was obtained for a total HRT of 184 h.

Vazquez et al (2006 b)

33. Pulp and Paper mill wastewater (COD = 650 mg/L)

The study was carried out in a subsurface flow wetland (30.7 m2). It was divided into eight cells each size of 3.2 x 1.2 x 0.8 m. The depth was 2.3 m.

Initial 15 months results, indicate that removal efficiencies for phenol were 60 % at 5 day HRT and 77 % at 3 day HRT. It was thought that the longer retention time might have caused oxygen and nutrient deficiencies.

Abira et al (2005)

81

2.8 COUPLED ADVANCED OXIDATION AND BIOLOGICAL

PROCESSES FOR WASTEWATER TREATMENT

Wastewaters from chemical, pharmaceutical, and dye industries

most often contain significant amount of non-biodegradable organic

compounds. The elimination of these non-biodegradable toxic contaminants is

required before biological treatment. Although the biological treatment of

wastewater is often the most economical alternative process when compared

to other treatment options, the ability of a compound to undergo biological

degradation depends on a variety of factors. Such factors include the

concentrations, chemical structures, and the biodegradability of the target

molecules. Characteristics of the wastewater, such as pH, alkalinity, or the

presence of an inhibitory compound matrix, could also pay an important role

in the biological degradation of pollutants. Although many organic molecules

are readily biodegradable, many other synthetic and naturally existing organic

molecules are biorecalcitrant i.e., resistant to biodegradation (Tabrizi and

Mehrvar 2004).

Depending on the nature of the pollutants and the level of

contaminants, detoxification might be difficult and/or expensive to achieve by

conventional biological methods. In such cases, biological processes alone are

not able to reach effluent standards for the discharge into municipal sewer or

into surface water, therefore a pre-treatment or post-treatment is required. The

choice of the correct combination system must be carried out considering

several factors, both technical (treatment efficiency, plant simplicity,

flexibility, etc.) and economical (capital and operating costs including reagent

and energy consumption , sludge and gas disposal, maintenance, etc.) aspects.

In several cases, specific experimental tests are required in order to assess

actual efficiency and proper treatment conditions. Moreover, advanced

oxidation processes (AOPs) such as UV, UV/H2O2, UV/O3, UV/H2O2/O3, and

82

UV/TiO2 have been used as an attractive alternative for the treatment of these

types of wastewaters. AOPs are technologies for the production of highly

reactive intermediates, mainly hydroxyl radicals (˙OH), which are able to

oxidize almost all organic pollutants. Advanced oxidation processes can

reduce pollutant concentrations, and some processes produce more oxidized

compounds, which are in most cases more easily biodegradable than the

former ones. Although AOPs are expensive to install and operate, they may

be unavoidable for the tertiary treatment of refractory organics present in

industrial effluents to allow safe discharge of industrial contaminants. Despite

the effectiveness of AOPs, there are several scenarios that make them

economically disadvantageous. Effective treatment of a particular industrial

wastewater may require a combination of AOPs and biological processes in

order to exploit their individual quantities and, thus, reach the desired quality

within reasonable economical limits.

On one hand, AOPs have shown their worthiness for toxic

compounds elimination in water and wastewater treatment, however, the total

mineralization through these processes is very expensive. On the other hand,

biological treatment is relatively cheap and reliable process but there are

substances, which are unable to deal with. A combination of both processes

would mean a cheaper option for total organic degradation from a toxic

wastewater or a wastewater containing refractory organics (Tabrizi and

Mehrvar 2004).

It has been shown that the combination of advanced oxidation and

biological processes has the following advantages:

i) Advanced oxidation and biological processes are

accompaniments of each other.

83

ii) AOP pre-treatment can protect the microorganisms from

inhibitory or toxic compounds.

iii) Decrease in chemical cost by using cost-effective biological

pre or post-treatment.

iv) Total residence time is flexible as a result of different choices

that are possible for photocatalytic and biological reactor

residence times in a constant efficiency.

v) Achieving total mineralization for the organics while

minimizing the total cost.

2.8.1 Strategy for Coupling AOP and Biological Treatments

As a general treatment strategy, four types of treatment for a

chemical compound are possible. In some cases only biological treatment

alone is sufficient to enhance the effluent quality. In the presence of some

refractory or toxic compounds in wastewater, chemical pre-treatment is

required. In case biological treatment is not sufficient for biodegradable

compounds, chemical post treatment is also necessary. In some rare cases,

combination of chemical and biological treatment in multi-stages is necessary.

A general strategy that can be used to develop a coupled advanced oxidation

and biological processes for the treatment of a certain wastewater, which

might contain non-biodegradable or toxic organics are, as a first step to avoid

utilization of high cost due to AOPs, it must be confirmed that whether the

wastewater contains recalcitrant or toxic organics. If the wastewater is

biodegradable, conventional biological reactors are used to treat the

wastewater.

If it is confirmed that wastewater contains recalcitrant or toxic

organic, it would be pre-treated by AOPs to modify the structure of pollutants

by transforming them into less toxic and easily biodegradable intermediates,

84

which are degraded in the subsequent biological reactor in a shorter time. This

method can also prove to be less expensive in comparison to the AOPs alone

and less time consuming compared to the biological process. Moreover, if the

wastewater from the final biological reactor has met the requirements, it will

leave the treatment plant, otherwise it has to go through the previous cycle

(Parra et al 2002). The information about the toxicity and the biodegradability

of wastewater treated by AOPs allows us to determine an optimum treatment

time in the AOPs reactor of the coupled system. The time should be the best

compromise between the efficiency of the chemical reactor and its cost. The

shorter reaction time avoids the high electrical cost of the treatment. At longer

photo-treatment time, the photochemical efficiency is improved by the

unnecessary photo-degradation of pollutants that are biological degraded.

However, the overall efficiency remains almost constant. This implies higher

energy consumption without beneficial effect, as about 60 % of the total

operational cost is electricity (Sarria et al 2003). However, if the reaction time

is too short, the intermediates remaining in the system could still have

toxicological or bio-recalcitrant effects. The four types of wastewater, which

have potential for increasing treatment efficiencies by coupled processes are

recalcitrant compounds, biodegradable wastes with small amounts of

recalcitrant compounds, inhibitory compounds and intermediate dead-end

products (Gogate and Aniruddha 2004 a).

2.8.1.1 Recalcitrant compounds

Chemical pre-treatment with oxidants is a capable method of

converting recalcitrant pollutants into easily biodegradable compounds.

However, biodegradable compounds, which are not oxidized by chemical

means, may slow down the whole process. This group consists of large

molecules such as polymers that cannot be degraded easily due to their large

size and not having enough reactive sites. Partial oxidation of COD is

85

favourable for subsequent biodegradation, whereas further mineralization can

reduce the efficiency of the biological treatment. If the following step is

biological treatment, pre-treatment favours the organic pollutants by an

increase in the maximum microbial specific growth rate and decrease in the

inhibitory effect by microorganisms. Isoproturon was removed by using

photocatalysis followed by fixed bed reactor successfully. The reaction time

with TiO2 was 60 min with complete removal of Isoproturon and 5% of

dissolved organic carbon remained in solution after the biological process.

But AOPs alone could only remove dissolved organic carbon by 20%

(Pulgarin et al 1999).

2.8.1.2 Biodegradable wastes with small amounts of recalcitrant

compounds

To achieve most effective cost effective treatment, only recalcitrant

compounds should remain after biological pre-treatment. The aerobic

treatment as a first step in treating black olive mills wastewater enhances the

later ozonation by removing the most biodegradable compounds. The post-

ozonation is capable of degrading the remaining non-biodegradable matters as

well as most of the phenolic compounds not removed before (Benitez et al

2003).

2.8.1.3 Inhibitory compounds

Using chemical treatment alone may not be an effective technique

to remove some highly toxic and stable compounds. The pretreatment is

necessary to modify the structure of contaminants by changing them into less

toxic and easily biodegradable by-products. This enhances the following

biological treatment to degrade the organics in a short period of time. Acute

toxicity measured as luminescence inhibition of Vibro fischeri was affected

86

by ozone as a first process. Therefore, the wastewater from ozonation can be

treated by biological process without the risk of having any inhibitory

compounds in the pre-treated tannery wastewater (Jochimsen and Jekel 1997).

Olive mill wastewater contains some organics that are toxic to some

methanogenic bacteria, and their elimination by pretreatment reduces the

toxicity of the subsequent aerobic treatment (Beltran et al 2001). Textile

wastewater was also pretreated by photochemical technology followed by an

activated sludge process successfully (Ledakowicz and Gonera 1999). The

pretreatment step decreased the inhibitory effect of wastewater and made it

suitable for biological treatment.

2.8.1.4 Intermediate dead-end products

Generally, to make sure that chemical pre-treatment can be utilized

as a first step followed by a biological process, it is important to obtain

information concerning the chemical nature of intermediates formed during

the pretreatment as well as to know the toxicity, biodegrability, and the

evolution of ions. Although toxicity may be decreased with chemical pre-

treatment, some compounds depending on the nature of the intermediates

produced and the chemical oxidants used may also increase the toxicity. It

was shown that after photo treatment, the solution resulting from the

degradation of metobromuron was not appropriate for the biological

treatment, as the intermediates were more toxic than the parent compounds. It

has also been shown that choosing the correct AOPs or biological treatment

had an important effect on the final efficiency of the wastewater. Moreover, to

have a much more effective treatment of wastewater, the first chemical

pretreatment has to be chosen properly, therefore it will facilitate the next

biological treatment. It was observed that different oxidants have different

influence on TOC removal of isoproturon and metobromuron as a function of

irradiation time (Parra et al 2000).

87

2.9 OVERVIEW OF WORK DONE ON COUPLED ADVANCED OXIDATION AND BIOLOGICAL TREATMENT

PROCESSES

Recently, the combination of advanced oxidation (AOP) and

biological treatment processes has been proposed. The coupling of

photocatalysis and a biological treatment (activated sludge) has been studied

by some authors, but literature is still scarce. From the Table 2.9 it was

observed that the combination of photocatalysis followed by activated sludge

was mainly used, showing that biodegradability of the effluents has always

been enhanced by the photocatalytic treatment with periods of irradiation

ranging from minutes to several hours. In most cases there was just one

photochemical reactor followed by biological reactor in series. However,

there are three cases in which there is a biological pre-treatment followed by a

chemical treatment step.

The AOP followed by the biological treatment (AOP-biological

treatment) could be justified if bio-recalcitrant compounds are easily

degradable by the AOP and the resulting intermediates are easily degradable

by the biological treatment. As the AOP-biological treatment, photocatalytic-

biological treatment of atrazine (Chan et al 2004), pulp mill bleaching

wastewater (Yeber et al 1999), 6-chlorovanillin (Yeber et al 2000), and azo

dyes and wool textile wastewater (Chun and Yizhong, 1999), photocatalytic/

photofenton–biological treatment of isoproturon (Sarria et al 2002),

ozonisation–biological treatment of phenolic acids (Amat et al 2003), and

UV-biological treatment of polycyclic aromatic hydrocarbons (Guieysse and

Viklund, 2005) and chlorophenols (Tamer et al 2006) have been reported. The

comparisons must be performed whether these combined treatments are better

than the AOP treatment only.

88

Table 2.9 The combined AOP and biological treatment of organic pollutants in water and wastewaters

Sl.No.

Chemicals/Wastewater treated

Order of Scheme

Experimental condition and type of equipment Important findings of the work Reference

1. Dyeing wastewater (COD = 900 -3000 mg/L)

Biological -AOP

An Intermittently Decanted Extended Aeration (IDEA) reactor was used as an aerobic process, which was constructed of Perspex with a height of 1m and 1D of 100 mm. The influent was pumped into reactor at a flow rate of 1L/d and retained for 3 days (HRT = 3 days) and SRT = 20 days. Recirculating batch plate type photocatalytic reactor was used and size of the reactor was 60 x 8 x 3 cm. The recirculated flow rate was in the range of 100 to 130 L/min. 4 g of TiO2 was coated on the Zeolite (0.8 to 1.8 mm) and 30 W/m2 blacklight lamp was used as light source.

The catalysed photooxidiation process can degrade those non-degradable organic substances in the effluent treated by the IDEA process and also decolorize the effluent completely. It was also found that some non-biodegradable organic substances can be converted to biodegradable forms by the sensitized photo-oxidation reaction. A bio-photoreactor system was designed to combine this photocatalytic reactor with the IDEA reactor for the treatment of dyeing wastewater. The performance of this combined bio-photoreactor system with and without recirculation was investigated and compared. The system with recycle has similar efficiency for decolourization and COD removal to that without recycle, but has a high capacity to eliminate the effects caused by a shock loading, and also the system can treat dyeing wastewater with a higher organic concentration.

Li and Zhao (1997)

89


Sl. No.


Order of Scheme


2. Cork processing wastewater (COD = 1900 mg/L)

AOP-Biological Biological -AOP

The continuous ozonation reactor was a 5000 cm3 glass bubble column operating at 200C. The aerobic process was performed in an activated sludge system. The HRT was kept as 96 h with constant temperature of 200 C. The MLVSS was maintained in the range of 1.2 – 1.8 g/L

The removal obtained in ozonation alone were 12-54 %, 65-81 % and 55-89 % for the COD, total phenolics, and absorbance at 254 nm respectively, the biodegradability varied from an initial 0.60 to final values of 0.93. The optimum hydraulic retention time and ozone partial pressure were 3 h and 3 Kpa, respectively. The COD removal obtained in aerobic treatment was between 13 % and 37 % for HRT between 24 and 96 h and the Contois model gave values of qmax = 0.14g COD /g VSS/ h and K1 = 22.6 g COD/g VSS . The sequential processes increased the substrate removal efficiencies in comparison with the individual processes. The enhancements were greater in the aerobic degradation-ozonation.

Benitez et al (2003)

3. 5-amino-6-methyl-2-benzimidazolone (AMBI) (0.2 mmol/L)

AOP-Biological

The coil photochemical reactor with 400 W, medium pressure Hg-lamp was used. The solution was recirculated in batch mode at 22 L/h through the illuminated part of the photoreactor. The fixed bed biological reactor consists of a column of 1 L capacity containing biolite colonized by activated sludge was used. The wastewater was recirculated at 6 L/h through the column.

The coupled reactor was operated in a semi-continuous made and an optimal pretreatment time of 300 min was found. The initial concentration of 0.2 mmol/L AMBI was completely removed and about 40 % of DOC was remained after 300 min in iron photoassited process. H2O2 and O2 were compared as electron acceptors in the pretreatment process, the higher biological efficiency was observed in the system with O2. In the coupled process complete removal of DOC was achieved. High concentrated 4000 mg/L real AMBI wastewater was also successfully degraded using solar photo-Fenton process with 40 KJ/L of accumulated solar energy, 80 % of AMBI was eliminated and a positive value of the AOS was reached.

Ledakowicz and Solecka (2000)

90


Sl. No.


Order of Scheme


4. Textile wastewater (COD = 2154 mg/L)

AOP-Biological

The ozonation process was carried out in a 1.5 dm3 stirred gas-liquid rector equipped with UV lamps of 150 W. The optimum ozone dose, hydrogen peroxide and UV irradiation were 100 mg/dm3, 1cm3/dm3 and 1 h respectively. Biodegradation was carried out at 250 C in 250 cm3 shaken flasks. The MLVSS was maintained in the range of 1-2 g/L

The preozonated wastewater with initial COD of 2154 mg/L leads to an increase of the maximum specific rate of substrate elimination from 0.04 to 0.06 mgO2/mgVSS/h and to a significant decrease of the Monod constant from 3378 to 759 mg O2/dm3. This change in kinetic parameters suggested a faster biodegradation and that pretreated pollutants are more available to biological oxidation. And, also the inhibitory action of microbial growth in the untreated wastewater decreased from 47 % to 10 % after O3/ H2O2/ UV pretreatment.

Ledakowicz and Ganera (1999)

6. Common industrial wastewater

AOP-Biological

The thin film reactor of 1.44 m length with 0.52 m width and fixed bed height of 0.1 cm was used in photocatalytic treatment. The up-flow anaerobic sludge blanket reactor was operated at 350 C with a working volume of 1.6 L capacity continuously. HRT was maintained at 38.5 hrs.

The photocatalytic treatment under sunlight for 40 hours reduced the colour and COD removal by 74 % and 62 % respectively, whereas the COD removal in biological treatment was only 18 % after 120 hrs. treatment. The photocatalytic treatment improved the BOD/COD ratio from 0.21 to 0.56. After photocatalytic and biological treatments the no. of peaks in GC analysis was reduced to 31, whereas the original sample has 121 peaks.

Pratapreddy et al (2002)

7. Isoproturon (0.2 mmol/L)

AOP-Biological

Coaxial photocatalytic reactor with TiO2

supported on glass rings and 36 W black lights was used. The reactor had a total volume of 1.5 L. The solution was fed continuously at 90 L/h. The fixed bed reactor consists of a column of 1 L capacity containing biolite colonized by activated sludge. The wastewater was recirculated at 6L/h through the column.

Isoproturon was completely eliminated and about 80% of DOC remained in solution after 60 min. of phototreatment. The biodegradability and toxicity tests performed during a photodegradation show that the solution becomes biocompatible. The optimum time to stop the phototreatment before feeding the treated water to the biological reactor was found to be 1 hour. In this coupled system, 100 % of the initial concentration of Isoproturon and 95 % of DOC were removed.

Parra et al (2002)

91


Sl. No.


Order of Scheme


8. Tannery wastewater AOP-Biological

Fenton’s process was carried out in a stirred batch reactor of 1 L capacity. Biological treatment was carried out in 1.8 L aerobic reactors and 0.6 L settling tank. HRT of 1 day was maintained. Biomass concentration up to 35 g MLVSS/L corresponding to an F/M ratio of about 0.01 to 0.02 d-

1was maintained.

The relationships of H2O2/Fe2+ and H2O2/COD were 9 and 4 respectively, reaching an organic matter removal of about 90%, subsequently; the oxidized effluent was fed to an ASP, The biological organic matter removal of the pretreatment wastewater ranged between 35% and 60 % for COD and from 60 % to 70 % for BOD. The sequential AOP pretreatment and biological treatment increased the overall COD removal up to 96% compared to 60 % without pretreatment.

Vidal et al (2003)

9. Herbicide/Pesticide wastewater (COD =540 mg/L)

AOP-Biological

Ozone treatment was performed in two stainless steel columns of volume 450 L each operated at dosage of 100-120 mg/L a HRT of 1.5 hrs. at ozone. The bioreactor of active volume about 200 L and the support media was Biolite. HRT was maintained in the range of 20 to 30 min.

The synergy between chemical and biological oxidation allowed removing herbicide and pesticides from starting concentrations in the order of 107-108 g/L to zero. Pre-ozonation improved biodegradability, so that the following biological treatment was above to contribute significantly to overall performance. The final polishing of the effluent was achieved by a post-ozonation step.

Mezzanotte et al (2003)

10. Ethylene di amin-e tetra acetic acid (EDTA) (2.5x10-3 M)

AOP-Biological

The photocatalytic studies were carried out in a 3L annular photoreactor equipped with a 30 W germicidal lamp. The titania 0.85 g was coaled on glass raschig rings. The activated sludge reactor was used for biological treatment.

About 50 % of EDTA degradation was reached after 150 min irradiation. Simultaneously a four-time increase in the biodegradability, measured as BOD5/COD ratio was observed. The activated sludge is not capable to degrade the complex EDTA but remove partially the COD and efficiently the BOD5 of the pretreated solution.

Carla et al (2003)

92


Sl. No.


Order of Scheme


11. 2,4-Dichlorophenol (100 mg/L)

AOP-Biological

The photochemical reactor consists of a Jacketed thermostatic 1.5 L equipped with three black blue lamp. Biodegradation reaction was completed in a 1.5 L aerobic sequencing batch reactor

Initial concentration of 65 mg/L H2O2 and 10 mg/L Fe (II) during 35 min irradiation time was sufficient for total 2, 4-Dichlorophenol removal. At these conditions, biodegradability was increased from 0 to 0.15. In a sequencing batch reactor, results showed that at cycle duration of 12 h, a TOC removal efficiency of 64 % was obtained.

Almomani et al (2003)

12. Landfill leachate, pulping wastewater, phenolic wastewater (COD = 2800, 2600, 4500 mg/L)

Biological -AOP

The bioreactor was an activated sludge reactor with aeration (7.5 L) and settling chamber (2.5 L). The HRT was 2 days, 3 days, 1 day for landfill leachate, pulping, and phenolic wastewater respectively. Ozone treatment was performed in a stainless steel column of volume 2.3 L.

The combined method, aerobic bio-oxidation with ozonation of recycled biologically treated wastewater, increases the degradation efficiency when compared to the conventional aerobic bio-oxidation method. COD removal for landfill leachate increased from 61 % to 95 % at an ozone dose of 30 mg/L, from 76 % to 89 % at 52 mg/L for pulping and from 60 % to 75 % at 60 mg/L for phenolic wastewaters. The activated sludge was not deteriorated and the specific oxygen uptake rate constant increased 15-20 % as a result of a small ozone dose of 2 mg/L. Thus combined process should be a prospective method in the purification of recalcitrant wastewaters.

Kamenev et al (2003)

13. 1-amino-8-napthol-3,6-disulfonic acid (H-acid) (50-150 mg/L)

AOP-Biological

Photocatalytic experiments were carried out in an immersion well photoreactor equipped with 400 W Medium-Pressure mercury vapour lamp. 1 g of TiO2 was used as catalyst. Biodegradation experiments were carried out in 250 mL. Erlenmeyer flasks.

In photocatalytic treatment about 70-90 % COD reduction was attained in 5 hours. The degradation obeys Pseudo-first-order. Photocatalytic pre-treatment of H-acid for 30 min ensures enhanced biodegradation of effluents. Complete removal of H-acid was achieved in the coupled system.

Mohanty et al (2005)

93


Sl. No.

Chemicals/ Wastewater

treated Order of Scheme


14. Pulp mill effluent AOP-Biological

Ozonation experiment carried out in a semi batch bubble column reactor made up of plexiglass of height 21 m and diameter 0.1 m. The ozone generator produced ozone consistently at a concentration of 0.11 mg/mL. Batch biological treatment was conducted in shake flasks. The initial MLSS of 1000 mg/L was maintained.

At an ozone dosage of 0.7-0.8 mg O3/mL wastewater, integrated treatment showed about 30% higher TOC mineralization compared to individual ozonation or biotreatment. Ozone treatment enhanced the biodegradability of the effluent (21 % COD reduction and 13 % BOD5 enhancement) allowing for a higher removal of pollutants. Ozonation at pH 11 was more effective that that at pH 4.5 in terms of generating more biodegradable compounds.

Bijan et al (2005)

15. Table olive manufacturing wastewater

AOP-Biological

Wet air oxidation experiments were conducted in a 0.6 L, stainless steel autoclave batch reactor. The system was pressurized with air to 1 pa. Copper (II) was used as catalyst. Aerobic biological experiments were carried out batch wise in a cylindrical 3L glass reactors immersed in a theromostatic bath. The MLVSS was maintained in the range of 1000-2000 mg/L.

COD removal in the range of 30-60% at 6 h of treatment has been achieved by using mild conditions (443-483 K and 3.0-7.0 Mpa of total pressure using air). The rate of the COD biodegradation was compared to the kinetics of the aerobic process without a previous chemical pre-oxidation. The calculated kinetic parameters showed the positive effect of the pre-treatment (maximum growth rate of 0.030.006 h-1 and 0.014 0.00014 h-1). Acclimation of microorganisms to oxygenated species formed in a chemical pre-oxidation step enhanced the efficiency of the biodegradation.

Rivas et al (2001)

16. p-nitrotolunee-ortho-sulfonic acid (p-NTS)

AOP-Biological

The coil photochemical reactor with 400 W, medium pressure Hg-lamp was used. The solution was recirculated in batch mode at 22L/h through the illuminated part of the photoreactor. The fixed bed biological reactor consists oef a column of 1L capacity containing biolite colonized by activated sludge was used. The wastewater was recirculated at 6 L/h through the column.

The photo Fenton treatment generates intermediates with very oxidized functional groups being non-toxic and as biodegradable in 30 min. operated in semi-continuous mode, it was found that the optimal time to stop the photochemical treatment before leading the treated water into biological reactor was 70 min. At this moment appropriate efficiency was reached for the best compromise between time and energy invested in both biological and overall treatment.

Pulgrain et al (1999)

94


Sl. No.


treated

Order of Scheme


17. 6-Chlorovanillin (186 ppm)

AOP-Biological

The photocatalytic degradation was carried out in a 1 L glass reactor, TiO2 was supported on glass Raschig rings. 125 W high pressure lamp was used for irradiation. Biodegradation was carried out with two 250 mL Erlenmeyer flasks with P. paucimobilis and B. cepacia.

About 50% degradation was obtained after 30 min with recirculation of the solution. Then, oxidized samples were submitted under aerobic conditions to bacterial degradation in P. paucimobilis (S37) and B. cepacia (PZK). Both selected aerobic bacteria degrade more efficiently the photocatalysed samples, A first-order kinetic was observed in both systems.

Yeber et al (2000)

19. Logyard run-off (COD = 8050 mg/L)

AOP-Biological Biological -AOP

Batch biological studies were carried out in a 15 L cylindrical jacketed Plexiglas reactor. Ozonation was carried out in a 3 L jacketed glass vessel equipped with ports for ozone inlet and outlet.

Batch biological treatment of logyard run off reduced BOD (1250), COD (8050) and tannin (1550) mg/L and lignin (TL) concentration of 99%, 80% and 90 % respectively. The efficiency of ozone as a pre and post – biological treatment stage was assessed. During ozone pretreatment TL concentration and acute toxicity were rapidly reduced by 70 % and 71 % respectively. Pre-ozontion reduced BOD and COD concentration by less than 10 %. Biological treated effluent was subjected to ozonation, it was observed that a reduction in COD and TL concentration, however no further improvement in toxicity was observed. Ozonation increased BOD by 38 % due to conversion of COD to BOD.

Zenaitis et al (2002)

20. 2,4,6-Trinitro-Toluene (TNT)

AOP-Biological

Photocatalytic experiments were performed in a batch recirculating annular ring photoreactor 15 W fluorescent lamp. TiO2 0.3g/L was used as catalyst. The fungal mineralization by Phanerochaete Chrysosporium in a 250 ml Erlenmeyes flask.

The extent of TNT mineralization was approximately 14 % by biological transformation alone and improved to approximately 32 % using the combined photocatalytic and fungal treatment. Six hours of photocatalytic pre-treatment resulted in the greatest extent of biological mineralization in the combined process.

Hess et al (1997)

95


Sl. No.


Order of Scheme


21. Herbicides (Diuron and Linuron)

AOP-Biological

Photo-Fenton experiments were conducted at 250.20 C in a cylindrical Pyrex thermostatic cell of 0.25 L capacity with 6 W black light was use as light source. Biological experiments were conducted in the 1.5 L capacity sequencing batch reactor. HRT 2 days and MLVSS in the range of 600-1000 mg/L was maintained.

The combined photo-Fenton and biological process can completely remove the herbicides from water, treated with Fe (II) 15.9 mg/L and H2O2 202 mg/L during 1 h of UV irradiation. And complete removal of TOC was achieved after biological treatment in a SBR.

Maria et al (2006)

22. Procion Red (250 mg/L)

AOP-Biological

Photo-Fenton experiments were carried out using a cylindrical pyrex thermostatic cell of 300 mL capacity with 6 W fluorescent lamp. Biological experiments were conducted in 1.5 L SBR. MLSS and MLVSS were maintained in the range of 2.85 g/L and 1 g/L respectively.

Best pre-treatment results were obtained in 60 min of photo-Fenton process with 10 mg/L Fe (II), 125 mg/L H2O2. At these conditions, BOD5/COD was increased from 0.1 to 0.35, with 39 % mineralization. Complete mineralization was achieved in SBR with 1 day HRT.

Garcia et al (2006)

23. Reactive azo dye (100 mg/L)

AOP-Biological

UV/H2O2 experiments were carried out in a cylindrical rector of 585 mL capacity with 60 W low-pressure mercury lamp. Biological experiments were conducted in 5 L cylindrical reactor. Plastic carriers were suspended in the bioreactor to faster growth of biofilms.

The UV/H2O2 treatment demonstrated that the dye and COD removal efficiency increased from 20 to 86 % and 3 to 39 % respectively, by modifying the most influential parameters like UV irradiation time from 10 to 30 min, initial H2O2 dosage from 100 to 500 mg/L and recirculation ratio from 600 to 0 %. Complete degradation was achieved in biological treatment with 1 day HRT.

Sudarjanto et al (2006)

96


Sl. No.


treated

Order of Scheme


24. Pesticides (100 mg/L)

AOP-Biological

Chemical oxidation process was carried out in a 5 L volume tank. The UV emitting device typically consisted of a stainless steel tube with coaxial Mercury vapour lamp (14 W). Ozone rate was of 1.2 g O3 /h. Biodegradation studies were conducted in 5 L stirred tank. COD: N: P ratio was maintained around 100:5:1

The O3 and O3/UV oxidation systems were able to reach 90 and 100 % removal of the pesticide respectively in a period of 210 min. The combined O3/UV system can reduce COD up to 20 % if the pH of the solution is above 4. Both pesticide degradation and COD removal in the combined O3/UV system follow Pseudo-first-order kinetics. More than 95% COD removal was achieved when treated wastewater by the O3/UV system was fed to the bioreactor.

Lafi and Alqodah (2006)

25. Hospital wastewater (COD=1350 mg/L)

AOP-Biological

Photo-Fenton experiments were carried out in a 1L cylindrical quartz reactor with low germicide lamp. Biological degradations studies were conduced in the activated sludge system, consists of cylindrical aeration glass vessels with a total volume of about 500 mL. COD:N:P ratio was maintained as 100:5:1 and the MLSS = 3000 mg/L was maintained

At the optimum conditions, a dosage ratio of COD:H2O2:Fe (II) at 1:4:0.1 and at pH 3, the biodegradability, in terms of BOD5/COD ratio increased form 0.3 to 0.52 and the oxidation degree, (AOS) levelled up from –1.14 to +1.58. In biological process, a maximum COD removal of about 30 % after a 72 h treatment time. In contrast, the COD of pre-treated wastewater by Photo-Fenton processes, attaining a COD removal of higher than 90% at the end of a 72 h treatment time.

Kajitvichya nukul and Nattapol (2006)

97


Sl. No.


treated

Order of Scheme


26. Phenol (100 mg/L) P-Nitrophenol (50 mg/L)

AOP-Biological

UV-irradiation tests were conducted in 25 aliquots of 6ml of mineral salt medium supplied with 50 mg/L P-NP and 100 mg/L phenol were transferred into 25 x 10ml glass tubes placed beside each other on a rocking shaker, 1g of TiO2 was added to each tube. Two 18W UV blue lamps were used for irradiation. Biological treatments were conducted in a glass flask of 35 mL capacity. 25 mL of irradiation solution with 1mL of acclimated consortia. The flasks were flushed with N2 gas to remove O2 and incubated for 14 d under continuous agitation and illumination.

Photocatalytic degradation of phenol and P-NP was well described by Pseudo-first-order kinetic with removal rate constants of 1.9 x 10-4 and 2.8 x10-4 min-1 respectively, when the pollutants were provided together and 5.7 x 10-4 and 9.7 x 10-4 min-1, respectively, when they are provided individually. Pre-treatment of the wastewater during 60 h removed 50 1 % and 62 2 % of phenol and P-NP but only 11 3 % of the initial COD. Subsequent biological treatment of the pre-treated samples removed the remaining contaminants and 81-83 % of COD.

Tamer et al (2007)

27. Phenol (200 mg/L)

Biological -AOP

In bioreactor 0.2 L of the adapted sludge was mixed with 0.8 L of solution A (250 mg/ L of phenol, 62.5 g/ L of chloride, and pH 6.5) and the system aerated for 3 h. Photocatalytic experiments were carried out in a 0.25 L thermostated cylindrical Pyrex reactor with medium pressure mercury vapour lamp of 250 W.

The degradation of phenol by a hybrid process (activated sludge + photocatalysis) in a high salinity medium (50 g /L of chloride) has been investigated. The sludge used from a municipal wastewater facility was adapted to the high salt concentrations prior to use. The initial phenol concentration was approximately 200 mg/ L and complete removal of phenol and a mineralization degree above 98% were achieved within 25 h of treatment (24 h of biological treatment and 1 h of photocatalysis).

Amour et al (2008)

98

In the biological treatment followed by the AOP (biological-AOP

treatment), biodegradable compounds are removed by the biological treatment

and non-biodegradable compounds are degraded by the AOP. However, in the

reported studies of biological–photocatalytic treatments of kraft pulp

bleaching wastewater (Balcioglu and Cecen 1999) and dyeing wastewater (Li

and Zhao 1997), the concentration of pollutants in biologically treated water

was high and not suitable for TiO2 photocatalytic degradation. Although

phenol is toxic, 200 mg/L and 1000 mg/L could be decomposed by biological

treatment in 40 h (Prieto et al 2002) and 340 h (Gonzalez et al 2001),

respectively. On the other hand, 40 and 80 mg/L phenol could be decomposed

by the photocatalytic treatment in 6 h (Augugliaro et al 1988) and 8 h

(Sivalingam et al 2004), respectively. However, it required a long time to

mineralize concentrated phenol with only biological or photocatalytic

treatment. Therefore, the combined biological-photocatalytic treatment of

phenolic wastewater has been proposed. Studies reported on coupled system

are very limited hence more focus is needed in terms of optimization of pre-

treatment stage and its consequent effect on the biological process. The

oxidant dosage must be so adjusted that it is completely utilized in the pre-

treatment stage only. It is also important to note that a detailed analysis of the

oxidation products must also be done, as sometimes it may happen that the

intermediates formed might be toxic towards the microorganisms.

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