Decolorization and Mineralization of Azo Dye - American Chemical

30 Mar 2009 - Tanks-in-Series Model with Backflow for Ozonation in a Semibatch Bubble Column. For modeling the dynamic behavior of ozonation in a...
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Ind. Eng. Chem. Res. 2009, 48, 7965–7975

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Dynamic Modeling and Simulation of Ozonation in a Semibatch Bubble Column Reactor: Decolorization and Mineralization of Azo Dye Orange II by Ozone Masahiro Tokumura, Takashi Katoh, Hiroki Ohata, and Yoshinori Kawase* Research Center for Biochemical and EnVironmental Engineering, Department of Applied Chemistry, Toyo UniVersity, Kawagoe, Saitama 350-8585, Japan

Ozone decolorization and mineralization of azo dye Orange II were examined. In order to elucidate the dynamics of the oxidation process by ozone, the rate-based model consisting of the rates of chemical reaction and gas-liquid mass transfer was developed. The overall enhancement factor was introduced to correct the underestimation of mass transfer driving force for fast reactions. In modeling, nonideal mixing characteristics in the gas and liquid phases were described using a tanks-in-series model with backflow. The parameters in the proposed mixing model were determined on the basis of a multiple circulation cell model and a velocity profile of recirculatory flow in a bubble column. Experiments for decolorization and mineralization of Orange II were performed to verify the capability of the proposed model in a semibatch bubble column. The decolorization completed rather quickly. However, the mineralization characterized by the total organic carbon (TOC) was incomplete and its rate was rather slow. It was found that the decolorization rate decreased with the increase in the initial Orange II concentration and increased with increasing ozone gas injection rate, ozone dosage, and solution pH. The proposed model could reasonably describe the present experimental results for the dynamic changes and the steady state of decolorization, mineralization, dissolved ozone concentration, and ozone concentration in the outlet gas. The proposed model might effectively handle the prediction of ozone oxidation process in a semibatch bubble column. 1. Introduction Ozone has been widely used in water and wastewater treatment because of its high oxidation potential.1,2 Synthetic dyes released to the environment in the form of effluents by textile industries cause ecological damages. Wastewater from these industries is highly colored and contains organics. Ozonation is important as a potential process for color removal of dyes.1,2 The chromophore groups with conjugated double bonds, which are responsible for color, can be broken down by ozone, and as a result, the color of the effluents decreases. Bubble columns have been commonly used as ozone contactors in industries because of their simple construction, operation, and maintenance.1-4 However, bubble columns are difficult to design because of the complexity of the flow pattern.5,6 The ozonation process involves mass transfer with chemical reactions, i.e., the ozone mass transfer from the gas to the liquid phase, the ozone decomposition, and the degradation of pollutants with ozone. An understanding of ozone absorption into water associated with ozone decomposition and pollutant degradation is essential for rational design and operation of an ozonation process in a bubble column. Therefore, the removal of organic pollutants by ozonation has been actively discussed.1,2 However, the published literature does not give sufficient understanding of ozonation because most studies have focused primarily on ozone oxidation kinetics and have been based on the ideal mixing in a bubble column. Most of available ozonation models taking account of nonideal mixing have been developed for steady state. Several studies are available on the dynamic process of pollutant oxidation by ozone in water.1,2 A dynamic model can provide valuable information for start-up and shutdown operations, optimal design, and operability studies besides a steady-state performance. For example, Gurol and Singer7 proposed a mathematical model for the dynamic process * To whom correspondence should be addressed. Tel.: +81-49-2391377. Fax: +81-49-231-1031. E-mai: [email protected].

of ozonation of phenol in a bubble column and Beltran et al.8 examined the kinetic model for the ozonation of atrazine in water. Uchiyama et al.9 studied dynamic performance of ozonation treatment for nonionic surfactants by applying a tanksin-series model to describe the mixing in the gas phase. In their mathematical models, however, the mixing in the liquid phase was assumed to be complete mixing. Because of the low solubility of ozone, the rate-limiting step in the ozonation process comes from the low gas-liquid mass transfer, which significantly depends on the mixing characteristics in a bubble column. Chen et al.10 applied the axial dispersion model to discuss dynamic behavior of ozonation in a nonideal mixing in a countercurrent bubble column. The axial dispersion model has been widely used to describe nonideal mixing in a bubble column.1,5 It should be noted, however, that the axial dispersion model can describe satisfactorily only mixing, which does not deviate largely from the plug flow. The tanks-in-series model is more realistic and advantageous as compared with the axial dispersion model. Mechanistic investigation, in which nonideal mixing in the gas and liquid phases is considered using the tanks-in-series concept besides the chemical reaction kinetics and mass transfer, has not appeared to model the dynamic process of ozonation in a bubble column. Information about the dynamic ozonation processes is still scarce. It should be noted, incidentally, that the dynamic modeling for the systems, in which only ozone mass transfer and ozone decomposition occur and the ozonation of pollutants is not included, has also been studied using an axial-dispersion model or a tanks-in-series model.1,11,12 Of course, an axial-dispersion model and a tanksin-series model have been also applied to discuss the steadystate performance of ozonation process in a bubble column.1,13,14 The objective of this study is to examine decolorization and mineralization of azo dyes by ozonation in a bubble column. Azo dyes are the most commonly used commercial dyes. By taking account of reaction kinetics in the liquid phase and gas-liquid mass transfer combined with the tanks-in-series

10.1021/ie802009j CCC: $40.75  2009 American Chemical Society Published on Web 03/30/2009

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Figure 1. Schematic of experimental setup.

model for nonideal mixing in the gas and liquid phases, the simulation model for ozonation in a semibatch bubble column is developed. The parameters describing nonideal mixing in the tanks-in-series model with backflow are the numbers of hypothetical tanks in the gas and liquid phases and the backflow rate through the hypothetical tanks in the liquid phase, and usually they are adjustable parameters. In this study, they are determined using a multiple circulation cell model proposed by Joshi15 and the central plume model for the liquid velocity profiles in a bubble column proposed by Ulbrecht et al.,16 respectively. Experiments were conducted to validate the proposed dynamic simulation model. 2. Experimental Section All experiments were conducted in a cylindrical bubble column of 0.095 m diameter and 1.00 m height illustrated in Figure 1. Its working volume was 5 L. The liquid height (HL) was 0.705 m. The reactor was operated in a semibatch manner with a known volume of liquid forming the batch through which gas was continuously sparged. The reaction temperature or temperature of the liquid phase in the bubble column reactor was controlled within 2 K using the air conditioner. Orange II (sodium salt form, C16H11N2NaO4S, formula weight ) 350.32) was used as the model azo dye. The initial concentrations of Orange II were varied in the range of 5-200 mg L-1. In order to react with Orange II, ozone was continuously bubbled into the semibatch bubble column. The ozone gas was generated from pure oxygen or dry air as the feed gas using a corona discharge ozone generator (ED-OG-R2+, Eco Design Co., Japan) and introduced at the bottom of the bubble column through a glass filter gas sparger. The inlet concentration and flow rate of ozone gas (YO3,in and QG) were varied from 10 to 60 mg L-1 and from 0.2 to 2.0 L min-1, respectively. The gas flow rate was measured with a rotameter. The exit gas from the top of the bubble column was fed to the catalytic ozone destruction unit and then released to the ambient atmosphere.

Decolorization and mineralization of Orange II were started by the injection of ozone gas to water including a known amount of Orange II. At the beginning of the ozonation process, therefore, the water in the bubble column contained no ozone. Orange II concentration (Cd), total organic carbon ([TOC]), dissolved ozone concentrations (CO3), and ozone concentrations in the inlet and outlet gas flows (YO3,in and YO3,out) were measured at definite time intervals. It is important to quantify the dissolved ozone and ozone off-gas for reliable design and operation of the ozone contactor. Liquid samples were manually withdrawn through the sampling port (z ) 0.7HL) into small volumetric flasks to analyze for Orange II concentration and TOC (Figure 1). The concentrations of Orange II in water were analyzed by measuring its absorption at 486 nm using a U-1100 Spectrometer (Hitachi Co., Japan).17,18 Although the intermediates formed during the mineralization were not identified in this study, some intermediates may have some color. It is probable, however, that the spectrum of most intermediates is probably 0.2906 (2) Turbulent flow regime:

ur /u0 ) 1.0 for (r/R) e 0.4004 ur r 2 r ) -1.373 - 0.980 ln + 0.323 for (r/R) > 0.4004 u0 R R (3)

()

where the liquid velocity in the plume or the center of the column, u0, is given as u0 ) (200µ/FDC)((FUGDC /µ)(UG2 /DCg)-0.25)0.92

(4)

The flow rate between the hypothetical tanks in the liquid phase, QL, is evaluated by the following integration: QL )

∫ ∫ 2π

0

R

aR

-urr dr dθ

(5)

For a semibatch bubble column, as shown in Figure 2, the rates of liquid upflow and liquid downflow corresponding to backflow in the proposed tanks-in-series model are equal to each other. The unsteady-state change in the gas phase ozone concentration for the jth tank ( j ) 2,..., NG), YO3,j, can be represented by9,21,23 dYO3,j dt

)

(

)( )

1 - φG P Q (Y - YO3,j) φG VL G O3,j-1 1 - φG KLaLε(HeYO3,j - CO3,k) (6) φG

(

)

where k ) roundup(j/P). The first and second terms on the righthand side of eq 6 account for gas flows from (j - 1)th stage and to (j + 1)th stage and mass transfer from the gas phase to the liquid phase, respectively. Ozone mass transfer has been recognized to be controlled within the liquid film immediately adjacent to the gas-liquid interface based on its low solubility. We introduced the overall enhancement factor ε on the basis of the enhancement factor and depletion factor used by Benbelkacem et al.1,24,25 and Chen et al.10 In mass transfer models, the driving force is given as the difference between the equilibrium concentration at the gas-liquid interface and the concentration in the bulk liquid. In the fast reaction regime, the reaction occurs close to the gas-liquid surface and all dissolved ozone is consumed within the liquid film. The actual driving force given by the slope of concentration profile at the gas-liquid surface might be larger than the driving force based on the concentration in the bulk liquid defined by (C*O3,i - CO3,i). This causes an underestimation of mass transfer driving force for ozone transfer from the gas phase to the liquid phase. In order to correct this underestimation of mass transfer driving force, therefore, the overall enhancement factor was introduced while the reaction is fast and occurs within the liquid film. The values of ε used in this study are given below. The rate of change in ozone gas concentration in the bottom stage ( j ) 1) may be written as dYO3,1 dt

Figure 2. Tanks-in-series model with backflow for a semibatch bubble column.

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)

(

)( )

1 - φG P Q (Y - YO3,1) φG VL G O3,in 1 - φG KLaLε(HeYO3,1 - CO3,1) (7) φG

(

)

For the outlet ozone gas concentration, i.e., the ozone concentration in the gas stream from the head space in a bubble column (VH ) 1.5 L), we have the following material balance equation:

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dYO3,out dt

)

QG (Y - YO3,out) VH O3,NG

(8)

By assuming that the gas in the head space is completely mixed, we can obtain eq 8 for the change in the outlet ozone gas concentration, which we measured in the experiment. The relationship between the dissolved ozone concentration, CO3, and the ozone gas concentration, YO3, is given by Henry’s law: CO3 ) HeYO3

(9)

On the basis of the values and correlations in the literature,1,26-28 we postulated that the Henry’s law constant for ozone, He, is 0.30. The rate of change in dissolved ozone for the ith tank (i ) 2,..., (NL - 1)) may be written as23 dCO3,i dt

where the term [TOC]eq is defined as the practical steady-state TOC value. The overall degradation process was assumed to obey second-order kinetics with respect to [TOC]. The orders of the mineralization kinetics were determined using the present experimental results for the change in TOC with time. The kinetics of the reaction of ozone with organic compound is typically second order, i.e., first order in ozone and first order in the compound.1,7 The following overall kinetic equation for degradation of Orange II involving the second-order reactions of the concentrations of Orange II and ozone is assumed: rd,i ) -kdCO3,iCd,i

The dissolved ozone concentration in the bottom stage (i ) 1) is given as dCO3,1 dt

) KLaLε(CO* 3,1 - CO3,1) + rO3,1 + RrTOC,1 + βrd,1 +

) KLaLε(CO* 3,i - CO3,i) + rO3,i + RrTOC,i +

βrd,i +

QL QL (CO3,i-1 - CO3,i) + (C - CO3,i) (VL /NL) (VL /NL) O3,i+1 (10)

dCO3,NL

∑

m)(i-1)P+1

HeYO3,m

(11)

Here, R and β represent the consumed ozone per TOC degradation rate and per Orange II degradation rate, respectively. The first, second, third, fourth, fifth, and sixth terms on the righthand side of eq 10 account for mass transfer from the gas phase to the liquid phase, the consumption of ozone by decomposition, the consumption of ozone used to degrade TOC, the consumption of ozone used to decompose Orange II, the liquid upflows from (i - 1)th stage and to (i + 1)th stage, and the liquid downflows from (i + 1)th stage and to (i - 1)th stage, respectively. Sotelo et al.29 examined the kinetics of ozone decomposition in water varying the temperature from 10 to 40 °C and solution pH from 2.5 to 9 and obtained the correlation of the ozone decomposition rate. It may be written as follows:

( -4T964 )C - 1.99 × 10 -10 130 exp( )[OH ] C T

rO3,i ) -3.26 × 105 exp

) KLaLε(CO* 3,NL - CO3,NL) + rO3,NL + RrTOC,NL + βrd,NL +

iP

CO* 3,i ) 1/P ·

14

O3,i

×

3/2 - 1/2 i O3,i

(12)

In this study, we used this correlation to represent the rate of ozone decomposition. Ozone reacts directly with pollutants and through radical chain reaction. The hydroxyl radicals, being powerful oxidants, are produced by the decomposition of ozone. The contributions of direct and indirect reactions to ozone oxidation process are difficult to measure, and at natural pH, both oxidant pathways are comparable.1,30,31 Benbelkacem et al.24 stated that direct reaction with crotonic acid is the main oxidation mechanism at pH 7. As described below, the experimental results indicate that the mineralization of Orange II was incomplete. By considering these results, we assumed the following overall kinetics for the mineralization of Orange II by ozone, rTOC,i ) -kTOCCO3,i([TOC]i- [TOC]eq)2

(13)

QL (C - CO3,1) (15) (VL /NL) O3,2

For the top stage (i ) NL), we have

dt

where

(14)

QL (C - CO3,NL) (16) (VL /NL) O3,NL-1

The unsteady-state change in TOC for the ith tank (i ) 2,..., (NL - 1)) can be represented as QL d[TOC]i ) ([TOC]i-1 - [TOC]i) + dt (VL /NL) QL ([TOC]i+1 - [TOC]i) + rTOC,i (17) (VL /NL) The first, second, and third terms on the right-hand side of eq 17 account for liquid upflows from (i - 1)th stage and to (i + 1)th stage, liquid downflows from (i + 1)th stage and to (i 1)th stage, and the degradation rate of TOC, respectively. We have the following equations for the changes in TOC in the bottom stage (i ) 1) and the top stage (i ) NL): QL QL d[TOC]1 ) [TOC]2 [TOC]1 + rTOC,1 dt (VL /NL) (VL /NL) (18) and d[TOC]NL dt

)

QL QL [TOC]NL-1 [TOC]NL + (VL /NL) (VL /NL) rTOC,NL (19)

respectively. By taking the unsteady-state material balance of Orange II for the ith tank (i ) 2,..., (NL - 1)), we have QL dCd,i QL ) (Cd,i-1 - Cd,i) + (C - Cd,i) + rd,i dt (VL /NL) (VL /NL) d,i+1 (20)

The first, second, and third terms on the right-hand side of eq 20 account for liquid upflows from (i - 1)th stage and to (i + 1)th stage, liquid downflows from (i + 1)th stage and to (i 1)th stage, and the degradation rate of Orange II, respectively.

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Figure 3. Typical experimental data for decolorization and mineralization of Orange II: Cd and [TOC] at z ) 0.7HL, CO3 at z ) 0.5HL. Lines are average and smooth curves of experimental data(Cd0 ) 50 mg L-1, YO3,in ) 30 mg L-1, QG ) 0.5 L min-1, and pH0 ) 7.3).

The decreases in Orange II concentration for the bottom stage (i ) 1) and the top stage (i ) NL) can be represented by QL QL dCd,1 )Cd,1 + C + rd,1 dt (VL /NL) (VL /NL) d,2

(21)

and dCd,NL dt

)

QL QL C C + rd,NL (VL /NL) d,NL-1 (VL /NL) d,NL

(22)

respectively. For the gas holdups, φG, the measured values in this study were used. The volumetric mass transfer coefficients for O2 transfer into water in the bubble column used in this study were measured by the dynamic method.32 By substituting the mass transfer data for O2 into the following correlation based on the penetration model, we evaluated KLaL for O3 transfer into water: (KLaL)O3 ) (DO3 /DO2)0.5(KLaL)O2

(23)

Values of diffusivity for O3 and O2 were calculated using the equation and data in the literature.9 By numerically solving the simultaneous ordinary differential equations presented, we can estimate the change in the concentration distributions of Orange II, TOC, and dissolved O3 in the liquid phase; ozone concentration profile in the gas phase; and outlet ozone gas concentration as a function of ozonation time. A fourth-order Runge-Kutta technique was applied to solve the ordinary differential equations. 4. Results and Discussions Figure 3 shows typical experimental results for the decolorization and mineralization of Orange II by ozonation. The experiment was performed at the initial Orange II concentration of 50 mg L-1, inlet ozone gas concentration of 30 mg L-1, and gas flow rate of 0.5 L min-1. When ozone gas was sparged into the column, the degradation of Orange II started immediately. The significant and rather fast decolorization occurred and was almost completed after 60 min. The solution turned from orange to colorless. Nitrogen-containing organic compounds having lower molecular weight might be created with the cleavage of the azo bond and further the intermediates were degraded.33,34 The degradation of Orange II followed by mineralization to

simpler organic intermediates. However, the mineralization rate was rather slow. After 90 min, only 55% of the initial TOC was removed and then the mineralization rate significantly slowed down. At 210 min ozonation time, the remaining TOC was 6.7 mg L-1 with TOC removal of 72.2%. Although the intermediates formed during the ozonation process were not identified in this study, this result might be due to intermediates such as phenol, β-naphthol, benzendicarboxylic acid, etc.,33 some of which were relatively difficult to be degraded by ozone oxidation. It can be seen from Figure 3 that the dissolved ozone concentration was practically kept to zero up to 95 min. Dissolved ozone was not observed until Orange II was completely decolorized. This observation might be attributed to the large consumption for ozone oxidation of Orange II and decomposition of ozone. All ozone absorbed from the gas phase to the liquid phase had been almost completely consumed to degrade Orange II and the intermediates up to 95 min. The ozone supply from the gas phase was smaller than the ozone consumption in the liquid phase. The oxidation of Orange II was a masstransfer-controlled reaction. After 95 min, the dissolved O3 concentration somewhat quickly increased and approached the steady-state dissolved ozone concentration. Only a part of the dissolved ozone was used to oxidize the intermediates, which are difficult to degrade only by ozone, and as a result, the dissolved ozone concentration increased to the steady-state concentration. After most of the Orange II was oxidized, mass transfer was no longer limiting, thereby reducing the consumption of dissolved ozone due to the oxidation of intermediates. Consequently, ozone dissolved into liquid phase accumulated, and then the concentration of dissolved ozone began to increase. Similar results for ozone concentration in the liquid phase were reported in the literature.35-37 Incidentally, the immediate increase in dissolved ozone concentration for ozonation in a semibatch reactor has also been reported.26,38 The practical applications of ozonation process are often restricted because of the high cost of ozone generation. Therefore, it is very important to understand the change in concentration of outlet ozone gas, which is uselessly exhausted. The ozone concentration change in the outlet gas was very complicated. It can be seen that the outlet ozone gas concentration change could be divided into six phases or regions. This change is similar to the results reported in the literature.10,39-41 It should be mentioned, however, that the continuous and

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Figure 4. Effect of the initial Orange II concentration on decolorization of Orange II: Cd/Cd0 at z ) 0.7HL. Lines represent the simulation results(YO3,in ) 30 mg L-1, QG ) 0.5 L min-1, and pH0 ) 6.5).

monotonous increase in the outlet ozone gas concentration has also been obtained.26,38 At the initial stage, the outlet ozone gas concentration rapidly increased (Phase-I) and then the initial increase in YO3,out somewhat slowed down (Phase-II). Phase-II was sometimes very short and indistinguishable. In Phase-I and -II, the decolorization was very effective. Subsequently, the ozone concentration in the outlet gas increased again up to 60 min (Phase-III). The termination of the decolorization caused a decrease in the enhancement of ozone transfer due to the ozonation in the liquid phase. However, ozone dissolved into the liquid was constantly consumed by the degradation of the intermediates formed during the decolorization process and the decomposition of ozone. The increase was somewhat decelerated for 60 < t < 90 min (Phase-IV) after the completion of Orange II decolorization. Since the consumption of ozone due to the reactions in the liquid phase was suppressed, the outlet ozone gas concentration again increased. The increase in YO3,out gradually accelerated up until 120 min (Phase-V). Then the increase in YO3,out accelerated because of the slowdown of the degradation of the intermediates, which are relatively difficult to degrade by ozone. It can be seen in Figure 4 that in Phase-V the decrease in TOC representing the mineralization of the intermediates was significantly suppressed. Successively the outlet ozone gas concentration approached the steady-state ozone gas concentration (Phase-VI). After about 200 min, the ozone concentration in the outlet gas and the dissolved ozone concentration reached steady-state values. It should be noted that the steady-state ozone gas concentration (YO3,out) was somewhat smaller than the inlet ozone gas concentration (YO3,in). After the steady state, in which the consumption of ozone by the oxidation of Orange II and the intermediates became insignificant, was achieved, the decomposition of ozone in the liquid phase was continuously taken into account. Therefore, we obtained YO3,in > YO3,out at steady state. The quick decolorization of Orange II, the fast degradation of intermediates, and the very slow mineralization correspond to Phase-I, Phase-II and Phase-IV, Phase-V, respectively. These observations presented in Figure 3 suggest that the rate of the dynamic ozonation cannot be expressed by a simple reaction-kinetic expression. During the ozonation process, the pH dropped from 7.3 to 5.5. This may be attributed to the formation of acid intermediates. After complete decolorization, the pH gradually increased from 5.5 to 6.7 as the dissolved ozone concentration increased. The degradation of acid intermediates may be responsible for this increase in solution pH. In Figure 4, effects of initial Orange II concentration, Cd0, on decolorization are depicted with the Orange II concentration range of 50-200 mg L-1. All the other operating parameters

Figure 5. Effect of inlet ozone gas concentration on decolorization of Orange II: Cd/Cd0 at z ) 0.7HL. Lines represent the simulation results (Cd0 ) 50 mg L-1, QG ) 0.5 L min-1, and pH0 ) 6.5).

were kept on constant. The color decreased exponentially with respect to time. When the initial concentration was 50 mg L-1, no dye was detected after 40 min. The decolorization process in the initial ozonation stage was nearly described by first-order kinetics with respect to the Orange II concentration. dCd ) -kd,appCd dt

(24)

The slope of a straight line for a plot of ln(Cd/Cd0) versus t is the pseudofirst-order reaction rate constant kd,app, which is useful to understand the influence of parameters such as pollutant loading, ozone dosage, and gas flow rate on the decolorization rate.1 It can be seen as shown in Figure 4 that the decolorization rate decreased with increasing initial Orange II concentration. Since the ozone dosage was kept constant, the decolorization rate decreased at higher Cd0. The effects of ozone dosage were examined by varying its amount from 10 to 60 mg L-1. All the other operating parameters were kept constant and were as follows: initial Orange II concentration of 50 mg L-1, ozone gas flow rate of 0.5 L min-1 and solution pH0 of 6.5. As seen in Figure 5, an increase in ozone dosage from 10 to 60 mg L-1 increased color removal efficiency from 46.0 to 99.7% at 20 min of ozonation time. With increasing ozone dosage, the overall decolorization rate constant for Orange II degradation, kd,app, almost linearly increased with the inlet ozone gas concentration. An increase in inlet ozone concentration enabled better mass transfer and, as a result, better oxidation of Orange II. Figure 6 illustrates effects of ozone gas flow rate on the decolorization. Experiments were conducted in the ozone gas flow rate range of 0.5-1.5 L min-1. All the other parameters were held constant. The inlet ozone gas concentration was kept constant at YO3,in ) 30 mg L-1. As the ozone gas flow rate increased, the decolorization rate increased. The applied ozone dose had a positive effect on decolorization. An increase in gas flow rate from 0.5 to 1.5 L min-1 increased color removal efficiency from 76.4 to 97.7% at 10 min of ozonation time. The reason could be that an increase in gas flow rate enhanced the mass transfer rate of ozone and then increased the supply of ozone used for decolorization and degradation of Orange II. In order to examine the validation of the proposed model for the ozonation of pollutants in water, first we discussed mass transfer of ozone and decomposition of ozone. Without addition of Orange II, we measured the concentrations of

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Figure 6. Effect of ozone gas flow rate on the decolorization of Orange II: Cd/Cd0 at z ) 0.7HL. Lines represent the simulation results (Cd0 ) 50 mg L-1, YO3,in ) 30 mg L-1, and pH0 ) 7.3).

Figure 7. Simulation of ozone transfer (without Orange II decolorization). Lines represent the simulation results: (a) QG ) 2.0 L min-1, QL ) 2.67 L min-1, KLaL ) 0.304 min-1, YO3,in ) 5.2 mg L-1, pH0 ) 7.9; (b) QG ) 0.2 L min-1, QL ) 0.926 L min-1, KLaL ) 0.0455 min-1, YO3,in ) 5.1 mg L-1, pH0 ) 7.8.

dissolved ozone and outlet ozone gas at QG ) 0.2 and 2.0 L min-1. Figure 7 compares the experimental data and the model predictions of changes in dissolved ozone concentrations and outlet ozone gas concentrations in a semibatch bubble column. The concentrations both in the liquid and gas phases continuously and monotonously increased and approached the steady-state concentrations. For QG ) 2.0 L min-1, the steady state was finally achieved after about 20 min. The steady-state outlet ozone gas concentration coincided with the inlet ozone gas concentration, YO3,in. This implies that the rate of mass transfer from the gas phase to the liquid phase was faster than the rate of ozone decomposition in the liquid phase, and then the theoretical equilibrium between the gas and liquid phases was achieved. The amount of ozone transferred from the gas phase to the liquid phase and decomposed was rather smaller as compared with the ozone in the gas phase at QG ) 2.0 L min-1. In fact, the steady-state dissolved ozone concentration of 1.43 mg L-1 was only slightly smaller as compared with the theoretical

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equilibrium ozone concentration in the liquid phase calculated using the ozone gas concentration and the Henry’s law constant. On the other hand, the concentrations for QG ) 0.2 L min-1 also continuously increased but could not achieve steady state within 60 min. A faster approach to the steady state could be achieved through a higher gas flow rate. This might be due to the lower mass transfer rate with the lower gas flow rate. It should be noted that the steady-state outlet ozone gas concentration (YO3,out) of 4.6 was somewhat less than the inlet ozone gas concentration. Furthermore, the steady-state dissolved ozone concentration was 0.86 mg L-1, which is 63% of the theoretical equilibrium ozone concentration in the liquid phase based on Henry’s law constant. After the steady state was accomplished, ozone was continuously transferred from the gas phase to the liquid phase, in which dissolved ozone was consumed by the decomposition. However, the dissolution rate of ozone into the liquid phase was rather smaller than the decomposition rate of ozone in the liquid phase, and as a result, the dissolved ozone concentration was rather smaller as compared with the theoretical equilibrium between the gas and liquid phases calculated using Henry’s law constant. The solid lines in the figure represent the simulation results that were obtained by solving the different equations for ozone gas and dissolved ozone, eqs 6-12, 15, and 16. It is seen from the figure that agreement between the experimental results and model predictions was satisfactory. This suggests that the values of volumetric mass transfer coefficient, gas holdup, and Henry’s constant and the correlation of ozone decomposition, eq 12, used in the proposed model are rational and applicable to discuss the ozonation of pollutants in water at around pH 7. The model predictions for decolorization and mineralization of Orange II are compared with the experimental results in Figure 8. Solid lines represent the model predictions. As can be seen in Figure 8 parts a and b, reasonable agreement between the model predictions and the experimental data for QG ) 2.0 and 0.2 L min-1 was obtained. By taking account of the overall enhancement factor ε, the proposed model, eqs 6 and 7, could predict the very complicated change in the outlet ozone gas concentration with time as well as the decolorization, mineralization, and dissolved ozone. It can be seen from Figure 8 that the continuous and monotonous decreases in Orange II and TOC concentrations, the increase in dissolved ozone concentration, which was first almost zero and then after the significant slowdown of mineralization rate increased and approached the steady-state concentration, and the increase in the outlet ozone gas concentration divided into six phases were reasonably simulated by the proposed rate-based model. In the modeling, as described above, the overall enhancement factor of ozone consumption ε was introduced to avoid the underestimation of the mass transfer driving force. We assumed its functional form as follows: ε ) (1 + εd + εTOC)

(25)

where b1Cd 1 + b2Cd b3([TOC] - [TOC]eq) ) 1 + b4([TOC] - [TOC]eq)

εd ) εTOC

(26)

The overall enhancement factors of ozone consumption were separately considered for Orange II degradation, εd, and TOC degradation, εTOC. Wu and Wang42 derived an empirical

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Figure 8. Dynamic simulation of ozonation of Orange II in the semibatch bubble column reactor. Lines represent the simulation results (Cd0 ) 51.9 mg L-1, YO3,in ) 5.7 mg L-1, pH0 ) 7.55). (a) QG ) 2.0 L min-1, QL ) 2.67 L min-1, KLaL ) 0.302 min-1, φG ) 0.013, kTOC ) 0.030 L2 (mg2 · min)-1, kd ) 11.0 L(mg · min)-1, R ) 9.4 mg (mg)-1, β ) 0.45 mg (mg)-1: (1) decolorization and mineralization of Orange II, Cd and [TOC] at z ) 0.7HL, CO3 at z ) 0.5HL; (2) enhancement factor and Hatta number at z ) 0.5HL. (b) QG ) 0.2 L min-1, QL ) 0.923 L min-1, KLaL ) 0.0457 min-1, φG ) 0.0013, kTOC ) 0.065 L2 (mg2 · min)-1, kd ) 13.0 L (mg · min)-1, R ) 4.6 mg (mg)-1, β ) 0.38 mg (mg)-1: (1) decolorization and mineralization of Orange II, Cd, and [TOC] at z ) 0.7HL, CO3 at z ) 0.5HL; (2) enhancement factor and Hatta number at z ) 0.5HL.

correlation of ε as a function of initial dye concentration, ozone dose, and temperature. Using the present experimental data, as well as the study of Wu and Wang,42 the values of the coefficients in the above equations, b1, b2, b3, and b4, were determined as 2.0, 0.5, 2.0, and 3.0 for QG ) 2.0 L min-1 and 4.2, 0.5, 0.1, and 0.5 for QG ) 0.2 L min-1, respectively. The variation of ε can be reflected by Hatta number for the ozonation of Orange II. The Hatta numbers for Orange II degradation and TOC degradation may be written as43 Had )

√βkdDO CO 3

3,i

reaction could be characterized by a fast reaction. In this regime, all reactions occur within the liquid film, and therefore, the actual driving force for ozone transfer from the gas phase to the liquid phase is significantly larger than the driving force based on the concentration in the bulk liquid. With continued ozonation, the reaction transitioned through a moderate then to a slow regime after ∼700 min. In the slow reaction regime, the reaction occurred within the bulk liquid and the enhancement factor was very close to 1. For QG ) 2.0 L min-1, the overall enhancement factor steeply

(27)

KL

and HaTOC )

√RkTOCDO CO ,i 3

KL

3

(28)

respectively. Mass-transfer-associated chemical reactions modeled in terms of the two-film theory can be briefly characterized by three different regimes, i.e., fast (Ha > 3), moderate (3 > Ha > 0.3), and slow (Ha < 0.3) regimes.1,35 The ozonation process started from the fast regime and changed to the moderate regime and finally reached the slow regime with time. In the case of QG ) 0.2 L min-1, as shown in Figure 9a, the value of (1 + ε) rather rapidly dropped from 9.3 to 2.2 after about 200 min and then very slowly decreased to 1. With time, the overall enhancement factor approached 1. The Hatta numbers for Orange II degradation and TOC degradation decreased from 2.9 to 0.3 and from 1.8 to 1.3 after about 200 min, respectively. During the first 200 min of ozonation, the Hatta numbers were >0.3 and the

Figure 9. Effect of solution pH on decolorization of Orange II: Cd/Cd0 at z ) 0.7HL. Solid and broken lines represent the simulation result for decolorization at pH0 ) 7.5 (natural pH) and the average and smooth curves of experimental data, respectively. (Cd0 ) 50.0 mg L-1, YO3,in ) 5.0 mg L-1, and QG ) 2.0 L min-1).

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decreased from 5.5 to 2.0 after ∼25 min. The corresponding decreases in Hatta numbers for Orange II degradation (Had) and TOC degradation (HaTOC) were from 4.5 to 0.25 and from 2.0 to 1.7 after 25 min, respectively. At the initial ozonation stage for t < 25 min, the ozonation could be characterized by a fast reaction that occurred within the liquid film. The overall enhancement factor significantly affected the simulation results at the initial ozonation stage. Similar changes in the enhancement factor with ozonation time were reported in the literature.10,24,25 Benbelkacem et al.24 obtained the quick drop of enhancement factor for ozonation of crotonic acid from 4.5 to 1.0 within 18 min. In Figures 4-6, the present model predictions for decolorization were compared with the experimental results obtained in the wide ranges of operating parameters and satisfactory agreement can be found. It is clear from the comparison of the results in Figures 7 and 8 that the time required to approach steady state increased with ozone consumption for Orange II degradation in the liquid phase. When the gas flow rate was 2.0 L min-1, the steady state only for ozone mass transfer was achieved after about 20 min but that for ozonation of Orange II was not achieved up to 1200 min. Solution pH plays a major role in the formation of OH radicals during ozonation. Experiments were conducted at pH0 values of 3.0, 7.5 (natural pH of the solution), and 11.0. Using sulfuric acid and sodium hydroxide, the initial solution pH was adjusted to 3.0 and 11.0, respectively. It is evident from Figure 9 that the decolorization efficiencies increased with increasing solution pH. The increase in decolorization at higher pH might be due to the enhancement of ozone decomposition by hydroxide ions generating OH radicals that can oxidize the organic compounds more efficiently as compared with molecular ozone.30,31 It is probable that, at high solution pH, substantial hydroxyl radical generation occurred and a major part of the Orange II degradation was via hydroxyl radical attack. It can be seen from the figure that satisfactory agreement was obtained between the model predictions and the experimental data for the natural solution pH. However, the simulation results for the results at pH0 3.0 and 11.0, which are not presented in the figure, could not fit the experimental results. In the present simulation model, the change in solution pH was taken into account only by the dependency of the solution pH on the ozone decomposition. The effect of solution pH on degradation rate of Orange II has to be considered. This is one of our future problems. It should be mentioned, incidentally, that eq 12 for the ozone decomposition is inapplicable in the pH range of pH > 9. It was found that the solution pH during the ozonation of Orange II decreased for all cases without controlling. The formation of organic acids probably reduced the solution pH.44 The decrease in solution pH or OH- ion concentration was suppressed by decreasing the initial pH, and the change in pH was insignificant at pH0 3.0. At low solution, the formation of organic acids was not sensitive to the change in solution pH. 5. Conclusions In order to simulate the dynamic process for ozonation of pollutant in a semibatch bubble column, we have developed the rate-based model in which the hydrodynamic behaviors, the ozone gas-liquid mass transfer, and the ozonation reaction kinetics are taken into account. Nonideal mixing characteristics in the gas and liquid phases were described using a tanks-in-series model with backflow, in which the

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parameters were determined on the basis of the multiple circulation cell model and the liquid velocity profile for recirculatory flow in a semibatch bubble column. The overall enhancement factor has been introduced to correct the underestimation of the mass transfer driving force for a fast reaction. We have conducted systematic experiments for Orange II ozonation to validate the proposed dynamic simulation model. Changes in the Orange II concentration, TOC concentration, dissolved ozone concentration, and outlet ozone gas concentration with ozonation time were measured in the semibatch bubble column. It was experimentally found that the decolorization completed rather quickly but the mineralization was incomplete and its rate was rather slow. It was also found that the decolorization rate decreased with the increase in the initial Orange II concentration and faster decolorization could be achieved through higher dosages of ozone. The decolorization efficiencies increased with increasing solution pH because the decomposition of ozone by hydroxide ions might be enhanced at high solution pH. The proposed model was validated by comparing its predictions with the present experimental data for ozonation of Orange II. The proposed model has proved to be adequate for describing the rather complicated dynamic performance of ozonation in a semibatch bubble column as well as the steady state. In particular, by introducing the overall enhancement factor, the proposed model could simulate satisfactorily the change in the outlet ozone gas concentration divided into six phases or regions. Although the coefficients in the correlation of the enhancement factor were somewhat arbitrarily determined using the experimental data in the present study, a more rigorous approach for the enhancement factor is desirable. One of our future research subjects is the study for the applicability of the model using the data in full-scale plants. Nomenclature a ) coefficient aL ) specific surface area based on liquid volume, m2 m-3 b1-b4 ) coefficients in eq 26, (mg L-1)-1 CO3 ) dissolved ozone concentration, mg L-1 CO3 ) dissolved ozone concentration at the gas-liquid interface, mg L-1 Cd ) Orange II concentration, mg L-1 DC ) column diameter, m D ) diffusion coefficient, m2 s-1 g ) gravitational acceleration, m s-2 HL ) liquid height, m Ha ) Hatta number defined by eqs 27 and 28 He ) Henry’s constant KL ) mass transfer coefficient, m(min)-1 kd ) reaction rate constant, L(mg min)-1 kd,app ) pseudofirst-order reaction constant in eq 24, min-1 kTOC ) reaction rate constant, L2(mg2 min)-1 NG ) number of hypothetical tanks for the gas phase NL ) number of hypothetical tanks for the liquid phase [OH-] ) OH- ion concentration, mg L-1 P ) ratio of the number of tanks for the gas phase to that for the liquid phase ()NG/NL) QG ) volumetric gas flow rate, L min-1 QL ) volumetric flow rate between the hypothetical tanks in the liquid phase, L min-1 R ) column radius, m r ) radial coordinate of column, m rd ) reaction rate for Orange II degradation, mg(L min)-1

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rTOC ) reaction rate for TOC, mg(L min)-1 rO3 ) reaction rate for ozone decomposition, mg(L min)-1 T ) temperature, K [TOC] ) total organic carbon concentration, mg L-1 t ) time, min UG ) superficial gas velocity, m s-1 ur ) radial liquid velocity distribution, m s-1 u0 ) liquid velocity at the center of the bubble column, m s-1 VG ) gas volume, L VH ) headspace volume, L VL ) liquid volume, L YO3 ) ozone concentration in the gas phase, mg L-1 z ) axial coordinate of column from bottom, m Greek Letters R ) consumed ozone per TOC degradation, mg(mg)-1 β ) consumed ozone per Orange II degradation, mg(mg)-1 ε ) overall enhancement factor of ozone mass transfer φG ) gas holdup µ ) liquid viscosity, Pa · s θ ) angular coordinate, rad F ) liquid density, kg m-3 Subscripts 0 ) initial i, j, k ) ith, jth, and kth tanks d ) dye in ) inlet O2 ) oxygen O3 ) ozone out ) outlet TOC ) total organic carbon

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ReceiVed for reView December 29, 2008 ReVised manuscript receiVed March 1, 2009 Accepted March 3, 2009 IE802009J