**Application of Mass Transfer Models in Environmental Engineering**

[23] Snyders, R., Music, D., Sigumonrong, D., Schelnberger, B., Jensen, J., Schneider, J. M.,

[26] Babić, B. M. S. K., Milonjic´, B. M. B., Polovina, M. J., Kaludierovi, B. V., Carbon 1999,

[27] Smičklas, I., Dimović, S., Šljivić, M., Plećaš, I., Lončar, B., Mitrić, M., J. Nucl. Mater.,

[29] Tóth, J. Adsorption Theory, Modeling, and Analysis, Marcel Dekker, Inc., USA, 2001.

[32] Barros, M.A.S.D., Zola A.S., Arroyo, P.A., Sousa-Aguiar, E.F., Tavares, C.R.G., Braz. J.

Appl. Phys. Letters, 2007, 90, 193902.

244 Mass Transfer - Advancement in Process Modelling

37, 477.

2010, 400, 15.

Chem. Eng., 2003, 20, 04, 413.

[24] Takeuchi, Y., Arai, H. J., Chem. Eng. Jpn., 1990, 23, 75.

[28] Pan, X., Wang, J., Zhang, D., Desalination, 2009, 249, 609.

[30] Markhan, E. C., Benton, A. F., J. Am. Chem. Soc., 1931, 53, 497.

[31] Jain, J. S., Snowyink, V. L., J. Water Pollut. Control Fed., 1973, 45, 2463.

[25] Park, J., Regalbuto, J. R., J. Colloid. Interface Sci., 1995, 175, 239.

Pen-Chi Chiang and Shu-Yuan Pan

Additional information is available at the end of the chapter

http://dx.doi.org/10.5772/60917

#### **Abstract**

Generally, unit operation processes that are used in environmental engineering are in‐ volved in interfacial reaction where mass transfer is an extremely essential component for system optimization. The purposes of this chapter were intended to provide the informa‐ tion of both theoretical model development and engineering practice for mass transfer of important processes in environmental engineering. Those processes include, but are not limited to, (1) ozonation (gas–liquid process), (2) ion exchange (liquid–solid process), (3) biological activated carbon (liquid–solid process), (4) chlorination (gas–liquid process), and (5) carbonation (gas–liquid–solid process).

**Keywords:** ozonation, ion exchanger, biological activated carbon, chlorination, carbo‐ naion, rotating packed bed

#### **1. Introduction**

#### **1.1. Ozonation process**

#### *1.1.1. Model development*

#### *1.1.1.1. Process chemistry*

Ozone process has been utilized in thousands of water treatment applications such as munic‐ ipal water (or wastewater) treatment plants, cooling towers, ultrapure water, marine aquaria, beverage industries, industrial process water, swimming pools, bottled-water plants, aqua‐ culture, and food processing. It can oxidize the micro-pollutants and reduce the amount of disinfection by-product (DBP) precursors formed in drinking water [1], such as organic peroxides, aliphatic aldehydes, hydrogen peroxides, mixed functional and saturated carbox‐

© 2015 The Author(s). Licensee InTech. This chapter is distributed under the terms of the Creative Commons Attribution License (http://creativecommons.org/licenses/by/3.0), which permits unrestricted use, distribution, and reproduction in any medium, provided the original work is properly cited.

ylic acids, etc. Some of the DBPs such as aldehydes and/or chlorinated aldehydes would be potential health hazards, leading to the increase of bacterial populations in distribution systems. Three oxygen molecules will combine together to form ozone, as shown in Equation (1); this is an endothermic reaction:

$$2\text{O}\_3 \leftrightarrow 3\text{O}\_{2'} \text{ } \Lambda \text{H}^0 \text{=} + 284.5 \text{ kJ/mol (at 1 atm)}\tag{1}$$

The gaseous ozone (O3) will decompose to oxygen gas (O2), as shown in Equation (2). In general, the decomposition rate increases as the temperature increases. The efficiency of ozone production decreases at a higher temperature with a fixed voltage input.

$$\rm O\_3 + O\_2 \rightarrow O\_2 + O\_2 + O\tag{2}$$

#### *1.1.2. Reaction kinetics and mass transfer*

Ozonation is a gas–liquid process under dynamic conditions; the reaction regime, mass transfer characteristics, and reaction kinetics are critical to system design and operation. As O3 dissolves in water, it will self-decompose and generate oxygen and hydroxyl-free radicals. In general, the solubility of dissolved O3 in ozonated water is merely 1 mg/L; higher dissolved O3 concentration will only happen in the environment with better mass transfer (e.g., mixing). The mass transfer rate is influenced by operating variables such as pH, temperature, agitation speed, and gas flow rate. In general, the ozone concentration generated from ozone generator usually is less than 14 % (v/v) that may restrict the mass transfer between gas and liquid phases. Therefore, residual ozone concentration in water treatment ranges from 0.1 to 1 mg/L. A higher dissolved ozone concentration may happen only if under better mass transfer conditions.

It is noted that the specific interfacial area is of great significance to the gas–liquid mass transfer. In ozonation process, the overall liquid-phase mass transfer coefficient (KLa) and equilibrium dissolved ozone concentration ([O3] \* ) can be obtained from the following mass balance equation [2]:

$$\frac{d\left[O\_{\text{3}}\right]}{dt} + r\_{O\_{\text{3}},sd} = K\_L a \left(\left[O\_{\text{3}}\right]^\ast - \left[O\_{\text{3}}\right]\right) \tag{3}$$

where *rO3,sd* is the self-decomposition rate of ozone, which can be determined by Yang [3] as Equation (4):

$$\begin{split} r\_{O\_3,sd} &= 5.117 \times 10^{10} \exp\left(-\frac{8835.25}{T}\right) \left[O\_3\right] + 9.207\\ &\times 10^{16} \exp\left(-\frac{8835.25}{T}\right) \left[OH^-\right]^{0.5} \left[O\_3\right] \end{split} \tag{4}$$

where *T* is the temperature (K). With the known [O3], t, and rO3,sd, both the KLa and [O3]\* can be determined through the use of a linear regression analysis under specific experimental conditions.

#### **1.2. Engineering practice**

ylic acids, etc. Some of the DBPs such as aldehydes and/or chlorinated aldehydes would be potential health hazards, leading to the increase of bacterial populations in distribution systems. Three oxygen molecules will combine together to form ozone, as shown in Equation

The gaseous ozone (O3) will decompose to oxygen gas (O2), as shown in Equation (2). In general, the decomposition rate increases as the temperature increases. The efficiency of ozone

Ozonation is a gas–liquid process under dynamic conditions; the reaction regime, mass transfer characteristics, and reaction kinetics are critical to system design and operation. As O3 dissolves in water, it will self-decompose and generate oxygen and hydroxyl-free radicals. In general, the solubility of dissolved O3 in ozonated water is merely 1 mg/L; higher dissolved O3 concentration will only happen in the environment with better mass transfer (e.g., mixing). The mass transfer rate is influenced by operating variables such as pH, temperature, agitation speed, and gas flow rate. In general, the ozone concentration generated from ozone generator usually is less than 14 % (v/v) that may restrict the mass transfer between gas and liquid phases. Therefore, residual ozone concentration in water treatment ranges from 0.1 to 1 mg/L. A higher dissolved ozone concentration may happen only if under better mass transfer conditions.

It is noted that the specific interfacial area is of great significance to the gas–liquid mass transfer. In ozonation process, the overall liquid-phase mass transfer coefficient (KLa) and equilibrium

> , 33 [ ] ([ ] [ ]) *O sd L d O r Ka O O*

where *rO3,sd* is the self-decomposition rate of ozone, which can be determined by Yang [3] as

8835.25 5.117 10 exp 9.207

*T*

æ ö

\*

3 3 \*

10 , 3

æ ö ´ -ç ÷

*r O*

16 0.5

=´ - + é ù ç ÷ë û è ø

*T*

è ø

8835.25 10 exp [ ] [ ]

*dt*

production decreases at a higher temperature with a fixed voltage input.

( ) <sup>0</sup>

3 2 2O 3O , H = + 284.5 kJ/mol at 1 atm D« (1)

O + O O + O + O 32 22 ® (2)

) can be obtained from the following mass balance

+= - (3)

(4)

3

*OH O*


(1); this is an endothermic reaction:

246 Mass Transfer - Advancement in Process Modelling

*1.1.2. Reaction kinetics and mass transfer*

dissolved ozone concentration ([O3]

3

*O sd*

equation [2]:

Equation (4):

#### *1.2.1. Determination of overall mass transfer coefficients*

Figure 1 presents a typical schematic diagram of ozonation process, where the ozone is produced by an ozone generator and able to reach 10 % or higher.

**Figure 1.** Schematic diagram of ozone experimental system (adapted from Chiang et al. [2])

Ouederni et al. [4] proposed the following equation to determine the ozone transfer in stirred reactors:

$$K\_L a = 2.57 \times 10^{-2} \times \text{(\textdegree C)}^{0.67} \tag{5}$$

where KLa is expressed in s-1, G is the gas flow rate expressed in 10-3 m3 /s, and n is the number of rotations per second. According to Sheffer and Esterson [5], the volumetric overall liquidphase mass transfer coefficient of ozone in tap water (KLa) was around 1.14 min-1. It was also observed that it would reach a higher value at a higher water flow rate (i.e., sufficient agitation). For instance, there was a large decrease in KLa to 0.25 min-1 without agitation.

In addition, combining the empirical relationship reported by Stankovic [6], the following equation can be derived:

$$K\_{\perp}a = h \cdot D\_v^{0.5} \cdot n^{0.5} \cdot G^{\times} \tag{6}$$

where *h* and *x* are constants, *a* is the specific interfacial area (cm-1), and *G* is the influent gas flow rate (L min-1).

#### *1.2.2. Determination of Fractional Ozone Absorption (FOA)*

To determine the optimal operation conditions for semi-batch ozonation process, the "frac‐ tional ozone absorption (FOA)," defined as the ratio of the ozone transferred to water to the ozone applied to the reactor [7], should be applied. The FOA value in semi-batch operation can be expressed as [2]

$$FOA\text{(\%)} = \frac{V \times K\_{\perp} a \times ([O\_3]^\ast - [O\_3])}{G \times P\_{O\_3} \times 1.98 \times 10^5} \times 10\tag{7}$$

where PO3 is the ozone partial pressure (kPa). The value of 1.98 x 105 is the factor which will transfer the term of denominator into the unit of mg/min.

An increase in temperature will increase the ozone diffusivity in water and enhance the liquidphase mass transfer, because the *KL* value is directly proportional to *Dv0.5* based on Higbie's penetration theory [8].

Figure 2 shows the relationships between the agitation speeds and influent gas flow rates under the specific conditions of KLa and FOA0. The results indicate that the FOA0 is mainly controlled by n (not by G). Both the values of *KLa* and *FOA0* (i.e., initial *FOA*) increase as the agitation speeds increase due to the breakdown of liquid film resistance and the enhancement of surface renewal rates. Therefore, the *KLa* increased with the increase of the influent gas flow rate since the specific interfacial surface area between gas and liquid phases increases. In addition, Figure 2 provides the baseline information of engineering to obtain the best achievable alternative among the various scenarios of *n*, *G*, and *KLa* values, in terms of the same level of FOA0. Furthermore, the desired FOA0 in the water treatment plants can be determined by the agitation speed, if the value of G is given.

The power consumption (Pg) should be taken into consideration for the most efficient design and operating conditions in practice. Pg can be calculated as follows [9]:

$$P\_g = \alpha \times N\_p \times n^3 \times d\_l^5 \times \rho \tag{8}$$

where *Pg* is the power consumption with gas dispersion (watt) and α is the ratio of *Pg* over *P0*, where *P0* is the power consumption in un-gassed liquid. Moreover, α is a function of superficial gas velocity so it is directly proportional to *G.*

**Figure 2.** Effect of agitation speeds and influent gas flow rates on KLa, initial fractional ozone absorption (FOA0), and power consumption. (Adapted from Chiang et al. [2])

*Np* is the power number:

In addition, combining the empirical relationship reported by Stankovic [6], the following

where *h* and *x* are constants, *a* is the specific interfacial area (cm-1), and *G* is the influent gas

To determine the optimal operation conditions for semi-batch ozonation process, the "frac‐ tional ozone absorption (FOA)," defined as the ratio of the ozone transferred to water to the ozone applied to the reactor [7], should be applied. The FOA value in semi-batch operation

> \* 3 3 5

1.98 10

3

*L O V Ka O O FOA G P*

([ ] [ ]) (%) <sup>10</sup>

where PO3 is the ozone partial pressure (kPa). The value of 1.98 x 105 is the factor which will

An increase in temperature will increase the ozone diffusivity in water and enhance the liquidphase mass transfer, because the *KL* value is directly proportional to *Dv0.5* based on Higbie's

Figure 2 shows the relationships between the agitation speeds and influent gas flow rates under the specific conditions of KLa and FOA0. The results indicate that the FOA0 is mainly controlled by n (not by G). Both the values of *KLa* and *FOA0* (i.e., initial *FOA*) increase as the agitation speeds increase due to the breakdown of liquid film resistance and the enhancement of surface renewal rates. Therefore, the *KLa* increased with the increase of the influent gas flow rate since the specific interfacial surface area between gas and liquid phases increases. In addition, Figure 2 provides the baseline information of engineering to obtain the best achievable alternative among the various scenarios of *n*, *G*, and *KLa* values, in terms of the same level of FOA0. Furthermore, the desired FOA0 in the water treatment plants can be determined by the

The power consumption (Pg) should be taken into consideration for the most efficient design

3 5

where *Pg* is the power consumption with gas dispersion (watt) and α is the ratio of *Pg* over *P0*, where *P0* is the power consumption in un-gassed liquid. Moreover, α is a function of superficial

 r

(8)

*g pi P Nnd* =´ ´ ´ ´

and operating conditions in practice. Pg can be calculated as follows [9]:

a

0.5 0.5 *<sup>x</sup> Ka h D n G L v* × × ×= (6)

´´ - <sup>=</sup> ´ ´´ ´ (7)

equation can be derived:

248 Mass Transfer - Advancement in Process Modelling

flow rate (L min-1).

can be expressed as [2]

penetration theory [8].

agitation speed, if the value of G is given.

gas velocity so it is directly proportional to *G.*

*1.2.2. Determination of Fractional Ozone Absorption (FOA)*

transfer the term of denominator into the unit of mg/min.

$$N\_p = \frac{P\_0}{n^3 \times d\_i^5 \times \rho} \tag{9}$$

Figure 2 can be used to determine *G* and *n* for given requirements of FOA0 and *Pg*, since the *Pg* value is influenced by both *n* and *G.* Under a constant value of *G*, the optimum operation conditions for semi-batch operation can be determined. For a standard six-bladed turbine with an *Np* value of 5.2 [9], the desired value of *α* can be determined corresponding to various G values.

#### *1.2.3. Integration of mass transfer and reaction kinetics*

Two important factors for determining the kinetic regime of ozonation are the enhancement factor (*E*) and Hatta number (*Ha*). The *E* factor is defined as the ratio between the apparent reaction and the maximum physical absorption rates:

$$E = \frac{N\_t}{K\_\perp a \left[O\_3\right]^\*}\tag{10}$$

where *Nt* is the actual ozone absorption rate (M s-1) and *KLa* is the overall mass transfer coefficient (*s*-1). The [O3]\* values can be determined from Henry's constant and the ozone partial pressure at the reactor outlet. On the other hand, the Ha number for O3 reaction is defined as

$$Ha = \sqrt{\frac{k\_1 \lbrack O\_3 \rbrack \choose k\_{O\_3, L}}{k\_L}} \tag{11}$$

DeCoursey [10] proposed a correlation between *E* and *Ha* for irreversible second-order reactions based on the Danckwerts penetration model, which takes the form as follows:

$$E = -\frac{Ha^2}{2\left(E\_o - 1\right)} + \sqrt{\frac{Ha^4}{4\left(E\_o - 1\right)^2} + \frac{E\_o \cdot Ha^2}{\left(E\_o - 1\right)} + 1} \tag{12}$$

where *E<sup>∞</sup>* is the infinite mass transfer enhancement factor. It is also noted that the expression is particularly accurate as the diffusivity coefficients of the reactants in the liquid are of the same order of magnitude.

Since the prediction model of *KLa* is applicable only to the deionized water, it is necessary to introduce the enhancement factor (*E*) to account for the effect of organic compounds in the real water sample on *KLa*. The *n* value (rps) for the fixed level of water quality parameters, applied ozone concentration [O3], and influent gas flow rate (Q) can be developed as

$$m = \left(\frac{Q\left[\text{O}\_3\right] + Vr\_{\text{O}\_3}}{1.20 \times EV\left(\frac{P\_{\text{O}\_3}}{0.38} - \left[\text{O}\_3\right]\right)} - 0.0225\right)^2 D\_v^{-1} \text{G}^{-236} \tag{13}$$

Therefore, the marginal level of ozone partial pressure (*PO3*) to achieve the highest removal of TOC can be determined through Equation (13). It is also noted that the mass transfer is successfully incorporated with reaction kinetics characterizations in Equation (13) for predict‐ ing the optimal *n* value (i.e., the most efficient agitation) to control the formation potential of organic compounds in the ozonation process.

Figure 3 gives the relationship between the agitation speed and applied ozone concentration under various retention times. For example, the TOC removal efficiency of 15 % and enhance‐ ment factor (E = 1) were held constant under the temperature of 25 °C, a pH value of 7, and an influent gas flow rate of 10 L/min. The higher the ozone partial pressure, the lower the agitation speed needed to achieve the same level of performance. In other words, it suggests that the ozone partial pressure (*PO3*) should slightly increase, instead of greatly enhancing the agitation power consumption, to obtain the same level of removal efficiency of TOC.

**Figure 3.** Simulation results of the relationship between agitation speed and ozone partial pressure in view of control‐ ling the formation of TOC. (Chiang et al. [2])

### **2. Ion Exchange (IE)**

3 \* 13 , 2 [ ] *O L L*

*<sup>k</sup>* <sup>=</sup> (11)


*kO D*

DeCoursey [10] proposed a correlation between *E* and *Ha* for irreversible second-order reactions based on the Danckwerts penetration model, which takes the form as follows:

2 4 2

where *E<sup>∞</sup>* is the infinite mass transfer enhancement factor. It is also noted that the expression is particularly accurate as the diffusivity coefficients of the reactants in the liquid are of the

Since the prediction model of *KLa* is applicable only to the deionized water, it is necessary to introduce the enhancement factor (*E*) to account for the effect of organic compounds in the real water sample on *KLa*. The *n* value (rps) for the fixed level of water quality parameters, applied

ozone concentration [O3], and influent gas flow rate (Q) can be developed as

3

0.38

power consumption, to obtain the same level of removal efficiency of TOC.

*O*

*P EV O*

*Q O Vr*

1.20

organic compounds in the ozonation process.

0.0225

ë û è ø è ø

ç ÷ é ù <sup>+</sup> ë û <sup>=</sup> - æ ö ´ - ç ÷ é ù

3

*O*

æ ö

*n D G*

3

Therefore, the marginal level of ozone partial pressure (*PO3*) to achieve the highest removal of TOC can be determined through Equation (13). It is also noted that the mass transfer is successfully incorporated with reaction kinetics characterizations in Equation (13) for predict‐ ing the optimal *n* value (i.e., the most efficient agitation) to control the formation potential of

Figure 3 gives the relationship between the agitation speed and applied ozone concentration under various retention times. For example, the TOC removal efficiency of 15 % and enhance‐ ment factor (E = 1) were held constant under the temperature of 25 °C, a pH value of 7, and an influent gas flow rate of 10 L/min. The higher the ozone partial pressure, the lower the agitation speed needed to achieve the same level of performance. In other words, it suggests that the ozone partial pressure (*PO3*) should slightly increase, instead of greatly enhancing the agitation

3 1 236

<sup>2</sup> 1

2

*v*


(13)

×

¥

 ¥

( ) ( ) ( )

¥

2 1 1 4 1 *Ha Ha E Ha <sup>E</sup> E E E*

=- + + + - -

*Ha*

¥

same order of magnitude.

250 Mass Transfer - Advancement in Process Modelling

Deionization is the process of removing the alts of electrolytes from water by ion exchange. "Deionization" is generally considered synonymous with "demineralization," by various authors, but since some of the ions exchanged are not necessarily of mineral origin, the term "deionization" appears more inclusive.

#### **2.1. Model development**

#### *2.1.1. Rate-determining step*

According to the Nernst static diffusion film theory, three mechanisms might be the ratedetermining step under appropriate conditions:


Recently, the principle of ion exchange kinetics has been investigated by applying either empirical rate equations or theoretical equations derived from fundamental diffusional mechanisms. In the latter approach, the rate-determining step could have been film diffusion, surface chemical reaction, or particle diffusion.

According to the kinetics of anion exchange on various resins, diffusion was the rate-control‐ ling process in all cases. Turse and Rieman III [11] reported that the rate of exchange was controlled by the chemical reaction if chelates were formed by cations and resins. Without formation of chelates, at low concentration, the rate was controlled by "film" diffusion and, on the other hand, at high concentration by "particle" diffusion. In other words, the ratedetermining step of the ion exchange process is always a combination of these two limiting steps of film diffusion and particle diffusion, except for the case of large organic ions (quater‐ nary ammonium ions) which are controlled by particle diffusion.

#### *2.1.2. Crank model*

The derivation of the Crank model is given below. For particle-diffusion controlled, the rate of exchange could be expressed as Fick's second law as follows:

$$\mathbf{D}\frac{\partial^2}{\partial r^2}(rc) = \frac{\partial}{\partial t}(rc) \tag{14}$$

With the initial condition:

$$\mathbf{r} \gets \mathbf{0}, \mathbf{0} \le \mathbf{r} \le \mathbf{a}, \mathbf{t} \le \mathbf{0} \tag{15}$$

and the boundary condition:

$$\left(\frac{\partial r}{\partial t}\right)\_{r=0} = 0, \ t > 0\tag{16}$$

Equation (14) can be easily solved by the Laplace transform. The *F (t)*, fraction of the electrolyte uptake by the resin, is expressed as [12]

$$F(t) = 1 - \sum\_{n=1}^{\sigma} \frac{6\sigma\left(\sigma + 1\right)}{9 + 9\sigma + \sigma^2 q\_n^{-2}} \exp\left[-\left(D\_p q\_n^{-2} + \frac{1}{\sigma^2}\right)\right] \tag{17}$$

where *Dp* 2 is the diffusion coefficient. *qn*s are the nonzero roots of Equation (17) as follows:

$$\tan\left(q\_n\right) = \frac{3q\_n}{3 + \sigma q\_n^{-2}}\tag{18}$$

In addition, *σ*is the ratio of the solute in the external solution to that in the resin particles at equilibrium, which can be determined by Equation (19):

$$
\sigma = \frac{3V}{4\pi^3 K} \tag{19}
$$

where *K* is the partition factor. Therefore, the diffusion coefficient can be determined with the aid of F (t) vs. Dpt/a2 , which is given by Crank (1957).

Huang and Tsai [13] have developed a more appropriate and general equation which com‐ bined film diffusion, particle diffusion, and chemical surface reaction on the exchange rate for a finite bath. Under a specified case, the rate equation (Table 1) using only isotopic exchange with particle resistance is consistent with the Crank model [12].


**Table 1.** Rate equation of isotopic ion exchange reaction with particle diffusion (P-mechanism), film diffusion (Fmechanism), and chemical reaction (C-mechanism) for various mathematical models

#### **2.2. Engineering practice**

controlled by the chemical reaction if chelates were formed by cations and resins. Without formation of chelates, at low concentration, the rate was controlled by "film" diffusion and, on the other hand, at high concentration by "particle" diffusion. In other words, the ratedetermining step of the ion exchange process is always a combination of these two limiting steps of film diffusion and particle diffusion, except for the case of large organic ions (quater‐

The derivation of the Crank model is given below. For particle-diffusion controlled, the rate

( ) ( ) <sup>2</sup> D <sup>2</sup> *rc rc r t*

0 0, 0 *r <sup>r</sup> <sup>t</sup> t* <sup>=</sup> æ ö ¶

( ) <sup>2</sup>

*n*

*m n F t D q*

tan

s s

s s

1

=

equilibrium, which can be determined by Equation (19):

¥ ç ÷ = >

Equation (14) can be easily solved by the Laplace transform. The *F (t)*, fraction of the electrolyte

6 ( 1) <sup>1</sup> <sup>1</sup> exp 9 9 *p n*

( ) <sup>2</sup> 3

*<sup>q</sup> <sup>q</sup>*

3

*q*

2 2 2

is the diffusion coefficient. *qn*s are the nonzero roots of Equation (17) as follows:

*n*

s

In addition, *σ*is the ratio of the solute in the external solution to that in the resin particles at

*n*

¶ ¶ <sup>=</sup> ¶ ¶ (14)

C=0,0 r a,t 0 ££ £ (15)

è ø ¶ (16)

s

*<sup>q</sup>* <sup>=</sup> <sup>+</sup> (18)

<sup>+</sup> é ù æ ö = - ê ú - + ç ÷ + + ë û è ø <sup>å</sup> (17)

nary ammonium ions) which are controlled by particle diffusion.

of exchange could be expressed as Fick's second law as follows:

*2.1.2. Crank model*

252 Mass Transfer - Advancement in Process Modelling

With the initial condition:

and the boundary condition:

where *Dp*

2

uptake by the resin, is expressed as [12]

#### *2.2.1. Determination of rate-determining steps*

In this chapter, the HPC and BHDAC are demonstrated as the representative quaternary ammonium ions because they are easy to prepare, highly soluble in water, and can be rapidly analyzed by use of UV. In addition, Dow-50-X8 was utilized as the resin for illustration. By plotting the *θ*<sup>t</sup> / *θ∞* for HPC and BHDAC on each Dow-50-X8 resin (i.e., 20–50, 50–100, and 100– 200 mesh), one can observe three straight lines with slopes of K (the parabolic rate constant). Subsequently, a liner correlation can be found between K and 1/D (the reciprocal of resin diameter), as shown in Figure 4. It suggests that the rate-determining step of these quaternary ammonium ions on the strong acid cation exchanger was particle-diffusion controlled.

**Figure 4.** Effect of particle size on parabolic rate constant (resin: Dow-50-X8)

#### *2.2.2. Determination of particle-diffusion coefficient*

The particle-diffusion coefficients of LPC, HPC, and BHDAC through the Amb-200, Amb-DPI, and Dow-50-X8 resins were determined and summarized in Table 2. As previously cited, the rate of exchange of HPC and BHDAC on various particle sizes on Dow-50-x8 resins was shown in Figure 4. It was observed that the Dow-50-X8 resin possessed the lowest particle-diffusion coefficient in comparison with the Amb-200 and Amb-DPI resins. This evidence strongly suggested that the macroporous resin had an advantage over the gel-structure resins for exchange of organic ions.


**Table 2.** Particle-diffusion coefficient of LPC, HPC, and BHDAC through the Amb-200, Amb-DPI, and Dow-50-X8 resins (20–50 mesh)

It was thus concluded that the Amb-200 (strong acid cation exchanger), Amb-DPI (weak acid cation exchanger), and XE-318 (chelating cation exchanger) resins were the ones to be selected as the representative cation exchangers for the resin selectivity test as described later.

### **3. Biological Activated Carbon (BAC)**

#### **3.1. Model development**

**Figure 4.** Effect of particle size on parabolic rate constant (resin: Dow-50-X8)

The particle-diffusion coefficients of LPC, HPC, and BHDAC through the Amb-200, Amb-DPI, and Dow-50-X8 resins were determined and summarized in Table 2. As previously cited, the rate of exchange of HPC and BHDAC on various particle sizes on Dow-50-x8 resins was shown in Figure 4. It was observed that the Dow-50-X8 resin possessed the lowest particle-diffusion coefficient in comparison with the Amb-200 and Amb-DPI resins. This evidence strongly suggested that the macroporous resin had an advantage over the gel-structure resins for

> **Amb-DPI (109 cm2 /sec)**

**Dow-50-X8 (109 cm2 /sec)**

*2.2.2. Determination of particle-diffusion coefficient*

254 Mass Transfer - Advancement in Process Modelling

**(109 cm2 /sec)**

LPC 3.03 2.35 -

HPC 2.43 2.12 0.85

BHDAC 0.90 1.70 0.57

**Table 2.** Particle-diffusion coefficient of LPC, HPC, and BHDAC through the Amb-200, Amb-DPI, and Dow-50-X8

exchange of organic ions.

resins (20–50 mesh)

**Quaternary ammonium ions Amb-200**

#### *3.1.1. Representative models of BAC*

The biological activated carbon (BAC) process contains adsorption and biodegradation mechanisms. The BAC process also exhibits lower regeneration cost and prolongs the life of granular activated carbon (GAC) beds. Therefore, it has been widely used in water and wastewater treatments.

A well-validated mathematical model can provide valuable information to evaluate and predict the performance of BAC process. Table 3 presents several representative models for BAC process. Chang and Rittmann [14] developed a model that the mass transfer of substrates diffusing through the biofilm, metabolized by microbes, and finally reaching the surface of GAC was illustrated and quantified. However, the limitation of the model is that it cannot be used under unsteady or plug flow conditions. Sakoda et al. (1996) developed a theoretical model for a BAC column considering the mechanisms including dispersion, convection, biodegradation, and adsorption. The assumption included that the substrate concentration on the interface between the biofilm and the GAC is identical to that in the bulk solution. Furthermore, Liang and Chiang [15] developed a non-steady-state numerical model to differentiate the adsorption and biodegradation quantities of a continuous BAC column including the mechanisms of adsorption, biodegradation, convection, and diffusion.



a A, adsorption; B, biodegradation; C, convection; D, dispersion.

b 1, bulk phase; 2, interface between bulk phase and biofilm; 3, biofilm; 4, interface between biofilm and GAC; 5, GAC. c Not analyzed in the article.

**Table 3.** Several representative models for biological activated carbon (BAC) process (modified from Liang et al. [16])

#### *3.1.2. Non-steady-state models for adsorption and biodegradation of BAC*

A numerical model can be developed to simulate both adsorption and biodegradation quantities of a BAC column under a non-steady-state condition. The governing equation based on mass balance of the substrate concentration in the liquid phase of the BAC column can be expressed as follows [16]:

$$\varepsilon \frac{\partial \mathbf{S}\_b}{\partial t} = D\_b \frac{\partial^2 \mathbf{S}\_b}{\partial \mathbf{x}^2} - v \frac{\partial \mathbf{S}\_b}{\partial \mathbf{x}} - \frac{\left(1 - \varepsilon\right)}{V\_g} \int\_0^{l\_f} \frac{k\_f X\_f \mathbf{S}\_f}{K\_f + \mathbf{S}\_f} 4\pi \left(r\_f + r\_g\right)^2 dr\_f - \left(1 - \varepsilon\right) \rho\_g \frac{\partial q\_s}{\partial t} - \varepsilon \left(\frac{k\_b X\_b \mathbf{S}\_b}{K\_b + \mathbf{S}\_b}\right) \tag{20}$$


Table 4 presents the parameter details of Equation (20):

**Reactor type**

Column A, B, C

Column A, B, C, D

Column A, B, C, D

Not analyzed in the article.

expressed as follows [16]:

a

b

c

e

Column B, C, D Nonsteady

**Mechanismsa Consideration of kinetics Mass**

Column A, B Uniform Monod Nonsteady n.a.c 1 Analytical

Monod Steady

Monod Nonsteady

Monod Steady

1, bulk phase; 2, interface between bulk phase and biofilm; 3, biofilm; 4, interface between biofilm and GAC; 5, GAC.

**Table 3.** Several representative models for biological activated carbon (BAC) process (modified from Liang et al. [16])

A numerical model can be developed to simulate both adsorption and biodegradation quantities of a BAC column under a non-steady-state condition. The governing equation based on mass balance of the substrate concentration in the liquid phase of the BAC column can be

**Biofilm amount** **Substrate in GAC**

Non-

Non-

Non-

Monod Monod Nonsteady n.a.c 1, 2, 3 Numerical

**Substrate in biofilm**

**Substrate in bulk phase**

256 Mass Transfer - Advancement in Process Modelling

Nonsteady, no biodegradati on

Nonsteady, no biodegradati on

Nonsteady, no biodegradati on, uniform Monod

A, adsorption; B, biodegradation; C, convection; D, dispersion.

2 0 1

e

*3.1.2. Non-steady-state models for adsorption and biodegradation of BAC*

( ) ( ) ( ) <sup>2</sup> <sup>2</sup>

*b fg f g*

4 1 *<sup>f</sup> <sup>L</sup> f ff b bb a b bb*

p

¶ ¶¶ - ¶ æ ö = -- + -- - ç ÷ ¶¶ + ¶ ¶ + è ø <sup>ò</sup> (20)

*S SS kXS q kXS D v r r dr t x x V KS t KS*

*g ff b b*

er

 e

**transport descriptionb**

equilibrium 1, 5 Analytical

equilibrium 1, 2, 3, 4, 5 Numerical

equilibrium 1, 2, 3, 4, 5 Numerical Liang and

**Solution method**

**Reference**

Walker and Weatherley [18]

Abumaizar, Smith and Kocher [19]

Hozalski and Bouwer [20]

Badriyha, Ravindran, Den and Pirbazari [21]

Chiang [15]

**Table 4.** Parameter details of non-steady-state models for adsorption and biodegradation of BAC

Two boundary conditions (BC) for the equation of dispersion advection reaction can be described as follows:

$$\text{BC}\,\mathbf{1}: \text{S}\_{b} = \text{S}\_{0} \quad , \quad \mathbf{x} = \mathbf{0}, \text{ and} \quad t \ge 0 \tag{21}$$

$$\text{BC.2}: \frac{\partial \mathbf{S}\_b}{\partial \mathbf{x}} = \mathbf{0} \Bigg|\_{\mathbf{x} = L\_c} \tag{22}$$

where Lc is the length of the BAC column (L). If the diffusion and reaction simultaneously occurred under a non-steady-state biofilm condition, the non-steady-state form of mass transfer and biodegradation reaction within biofilm, based on Fick's law and Monod equation, can be expressed as

$$\frac{\partial \mathbf{S}\_f}{\partial t} = \mathbf{D}\_f \frac{\partial^2 \mathbf{S}\_f}{\partial r\_f^2} - \frac{k\_f \mathbf{S}\_f}{\mathbf{K}\_s + \mathbf{S}\_f} \mathbf{X}\_f \qquad 0 \le r\_f \le \mathbf{L}\_f \tag{23}$$

where Df (L2 /T) is the diffusivity within the biofilm. A diffusion layer exists between the bulk solution and the biofilm, and the substrate concentration profile can be solved according to Rittmann and McCarty [22]. Therefore, the boundary condition can be simplified as

$$\left. \mathcal{S}\_f \right|\_{r\_f=0} = k\_{bf} \mathcal{S}\_b \tag{24}$$

where kbf is a factor to estimate the concentration reduction within the diffusion layer. Langmuir isotherm was used for the calculation of the boundary concentration of the biofilm near the GAC side, which was derived from the solid-phase concentration of the adsorbates. For single-component adsorbate, the surface concentration (qs, M/M) can be expressed as follows:

$$q\_s = \frac{q\_0 K\_L \mathbf{C}\_a}{1 + K\_L \mathbf{C}\_a} \tag{25}$$

where *q0* (M/M) is the unit layer adsorption capacity, KL (L3 /M) is the Langmuir coefficient, and Ca (M/L3 ) is the concentration on the boundary of the biofilm. It is assumed that the concentration flux from the biofilm should be identical to the substrate absorbed on a GAC granule, as Equation (26):

$$\left. \mathbf{4} \pi r\_{\clubsuit}^2 \cdot \mathbf{D}\_f \cdot \frac{\partial \mathbf{S}\_f}{\partial r\_f} \right|\_{r\_f = \mathbf{L}\_f} = \frac{\partial q\_s}{\partial t} \, \mathbf{m}\_g \tag{26}$$

where mg (M) is the mass of a GAC granule. In a control volume, the average biodegradation rate can be derived by integrating the Monod reaction expression and the amount of biofilm volume:

$$\left[\int\_{0}^{L\_{f}} \frac{k\_{f}S\_{f}}{K\_{f} + S\_{f}} X\_{f} 4\pi \left(r\_{g} + r\_{f}\right)^{2} dr\_{f}\right] \times N\_{g} \qquad 0 \le r\_{f} \le L\_{f} \tag{27}$$

$$N\_g = \frac{\Delta V \left(1 - \varepsilon\right)}{V\_g} \tag{28}$$

where ΔV is a control volume unit of the BAC bed (L3 ) and Ng is the number of BAC granules in a control volume.

#### *3.1.3. Estimation of biofilm thickness*

2

*S S kS*

where Df

follows:

and Ca (M/L3

volume:

in a control volume.

granule, as Equation (26):

(L2

258 Mass Transfer - Advancement in Process Modelling

*t KS r*

<sup>2</sup><sup>0</sup> *<sup>f</sup> f ff*

*f s f*

¶ ¶ = - £ £

Rittmann and McCarty [22]. Therefore, the boundary condition can be simplified as

0 *<sup>f</sup> <sup>f</sup> bf b <sup>r</sup> S kS* <sup>=</sup>

> 0 1

*s*

where *q0* (M/M) is the unit layer adsorption capacity, KL (L3

<sup>2</sup> 4

p

*f f*

*k S*

*K S*

where ΔV is a control volume unit of the BAC bed (L3

*L a*

*L a*

concentration flux from the biofilm should be identical to the substrate absorbed on a GAC

*f a g f g <sup>f</sup> r L <sup>S</sup> <sup>q</sup> r D <sup>m</sup> r t*

*f f*

where mg (M) is the mass of a GAC granule. In a control volume, the average biodegradation rate can be derived by integrating the Monod reaction expression and the amount of biofilm

*f gf f g f f*

ò (27)

*X r r dr N r L*

( ) 1

*g*

*V* D e

=

( )<sup>2</sup> <sup>0</sup> 4 0 *<sup>f</sup> <sup>L</sup> f f*

ê ú + ´ ££

*V*

p

*g*

*N*

é ù

<sup>+</sup> ë û

) is the concentration on the boundary of the biofilm. It is assumed that the

*f f ff*

solution and the biofilm, and the substrate concentration profile can be solved according to

where kbf is a factor to estimate the concentration reduction within the diffusion layer. Langmuir isotherm was used for the calculation of the boundary concentration of the biofilm near the GAC side, which was derived from the solid-phase concentration of the adsorbates. For single-component adsorbate, the surface concentration (qs, M/M) can be expressed as

/T) is the diffusivity within the biofilm. A diffusion layer exists between the bulk

¶ + ¶ (23)

= (24)

*qKC <sup>q</sup> K C* <sup>=</sup> <sup>+</sup> (25)

¶ ¶ × = ¶ ¶ (26)

= (28)

) and Ng is the number of BAC granules

/M) is the Langmuir coefficient,

*D X rL*

On the other hand, the thickness of biofilm will increase due to growth and decrease by the shear of water and the self-decay of bacteria. As a result, the biofilm thickness can be described as

$$\frac{\partial \mathcal{L}\_f}{\partial t} = \frac{Y \int\_0^{L\_f} \frac{kS\_f}{K\_s + S\_f} X\_f 4\pi \left(r\_s + r\_f\right)^2 dr\_f}{A\_f X\_f} \cdot b\_{\rm tot} L\_{f'} \quad 0 \le r\_f \le L\_f \tag{29}$$

where Y is the yield coefficient of biomass (CFU/M) and Af is the surface area of a BAC granule (L2 ).

In addition, the governing equation for the bacterial density in bulk solution can be simplified as an advection–reaction form:

$$\frac{\partial \mathbf{X}\_b}{\partial t} = -\nu \frac{\partial \mathbf{X}\_s}{\partial \mathbf{x}} + \frac{\mathbf{Y}k\_b \mathbf{S}\_b \mathbf{X}\_b}{K\_b + \mathbf{S}\_b} \tag{30}$$

In this situation, the BCs of Equation (30) can be expressed as follows:

$$\text{BC.1}: X\_b = \text{S}\_0 \quad \text{x} = \text{O}, \ t \ge 0 \tag{31}$$

$$\text{BC2} : \frac{\partial S\_{\text{b}}}{\partial \mathfrak{x}} = 0 \Bigg|\_{\mathfrak{x} = l\_{\text{v}}} \tag{32}$$

The governing equation is a second-order partial-differential equation, which can be approx‐ imately solved by the Crank–Nicolson finite differential and Crout factorization methods. Therefore, the developed model can quantify the amounts of adsorption and biodegradation. In addition, the significances of each parameter can be determined throughout the results of sensitivity analysis.

#### **3.2. Engineering practice**

#### *3.2.1. Dimensionless analysis*

The mass transfer coefficient can be determined by the dimensionless groups of Reynolds (Re), Sherwood (Sh), and Schmidt (Sc) numbers. Since these dimensionless groups are developed for process scale-up, different reactors can be compared to each other through the dimension‐ less analysis. The Reynolds number represents the ratio between inertial force and viscous force, and the Sherwood number represents the ratio between liquid-film transfer and biofilm diffusive transfer. The NSh can be correlated with the dimensionless groups as follows (Wakao and Funazkri, 1978):

$$N\_{\rm Sh} = 2 + 1.1 \times N\_{\rm Re}^{-1/2} N\_{\rm Sc}^{-1/3} \tag{33}$$

Several key operational parameters reasonably dominating the performance of a BAC column include the superficial influent velocity and the particle size of packing. As a result, the effects of particle size on adsorption and biodegradation can be simulated by the dimensional analysis.

#### *3.2.2. Liquid-film mass transfer*

Figure 5 shows the relationship between the Stanton number (NSt) and the liquid-phase mass transfer coefficient (kbf) reported in the literature. It is observed that the Stanton number (i.e., the ratio of the liquid-film transfer to the bulk transfer) exhibits an inverse proportion to the particle size (*dp*) and the superficial velocity (*vs*). In other words, both larger particle sizes and higher superficial velocities will lead to a relatively lower performance in liquid-film mass transfer (*NSt*). Therefore, the mass transfer from bulk solution into the biofilm can be enhanced with the increase of the Stanton number.

**Figure 5.** Correlation between the Stanton number (NSt) and the liquid-film mass transfer coefficient (kbf). *Dp* (cm) is the particle size, *vs* (cm/min) is the superficial velocity, and *Df* (10-6 cm2 /s) is the diffusivity. Available data was from Liang et al. [16]) and Badriyha et al. [21])

### **4. Chlorination process**

#### **4.1. Model development**

force, and the Sherwood number represents the ratio between liquid-film transfer and biofilm diffusive transfer. The NSh can be correlated with the dimensionless groups as follows (Wakao

Several key operational parameters reasonably dominating the performance of a BAC column include the superficial influent velocity and the particle size of packing. As a result, the effects of particle size on adsorption and biodegradation can be simulated by the dimensional

Figure 5 shows the relationship between the Stanton number (NSt) and the liquid-phase mass transfer coefficient (kbf) reported in the literature. It is observed that the Stanton number (i.e., the ratio of the liquid-film transfer to the bulk transfer) exhibits an inverse proportion to the particle size (*dp*) and the superficial velocity (*vs*). In other words, both larger particle sizes and higher superficial velocities will lead to a relatively lower performance in liquid-film mass transfer (*NSt*). Therefore, the mass transfer from bulk solution into the biofilm can be enhanced

**Figure 5.** Correlation between the Stanton number (NSt) and the liquid-film mass transfer coefficient (kbf). *Dp* (cm) is the

(10-6 cm2

/s) is the diffusivity. Available data was from Liang

1/2 1/3 2 1.1 *N NN Sh Re Sc* =+ ´ (33)

and Funazkri, 1978):

260 Mass Transfer - Advancement in Process Modelling

*3.2.2. Liquid-film mass transfer*

with the increase of the Stanton number.

particle size, *vs* (cm/min) is the superficial velocity, and *Df*

et al. [16]) and Badriyha et al. [21])

analysis.

Chlorination (using chlorine as disinfectant) is the common process used in daily water disinfection because of its high efficiency with a relatively low cost. However, as aforemen‐ tioned in the Ozonation chapter, the disinfection by-products (DBP), specifically trihalome‐ thane (THM) including chloroform, bromodichloromethane, dibromochloromethane, and bromoform, will also be formed during the chlorination process. Control of the formation potential of THM during chlorination is thus an important target for the supply of drinking water. THMs can be formed in the presence of specific precursors such as humic substances and methyl ketones [23]. Several good engineering practices to prevent THM formation include (1) decrease of chlorine dose, (2) change of the point of chlorination, (3) substitution with an alternative disinfectant, and (4) reduction of precursors in solution by GAC or enhanced coagulation. In addition, several factors including precursor concentration (as represented by total organic carbon concentration), chlorine dose, bromide concentration, reaction temperature, pH of solution, contact time, and reactor configuration (length-to-width ratio and mixing conditions) will affect the formation of THMs. To predict the performance of disinfection efficiency and kinetics, and/or THM formation potential, numerous models such as dispersion model have been applied in chlorine process, as illustrated in the following context.

#### *4.1.1. Dispersion model*

In general, the chlorination process is applied in either a rectangular basin with baffles (rectangular-long and narrow channel) or a circular basin (circular-annular ring surrounding). The hydrology in a rectangular basin with baffles is the most advantageous because of a plugflow-like reactor. It can be examined by the exit age distribution (*Eθ*) curve (Equation (34)) of the basin reactor, assuming the independence of fluid element and agitation with respect to radial position [24]:

$$E\_{\theta} = \frac{1}{\sqrt{4\pi\theta d}} \exp\frac{-\left(1-\theta\right)^{2}}{4\theta d} \tag{34}$$

where *d* is the dimensionless diffusion coefficient representing the degree of axial mixing. The *d* value approaching zero represents the reactor behaving as an ideal plug-flow type, otherwise (approaches infinity) representing a mixed-flow type. The *Eθ* in Equation (34) is a function of *θ* which follows a Gaussian distribution.

For a baffled rectangular basin with a total flow length of L and a channel width of W, the dimensionless diffusion coefficient (*d*) in open-channel flow, taking the nonideal nature of the reactor into consideration, could be expressed by Equation (35) [25]:

$$d = \frac{0.14k}{L/W} \tag{35}$$

where k is a nonideality coefficient of the reactor.

#### *4.1.2. Disinfection kinetics*

According to Collins and Selleck [26], the die-off of microorganisms in a batch reactor due to disinfection can be expressed by Equations (36) and (37):

$$
\left[\frac{N}{N\_0}\right]\_{\text{batch}} = 1,\quad \text{for } \theta < \frac{b}{CT} \tag{36}
$$

$$
\left[\frac{N}{N\_0}\right]\_{\text{bach}} = \left(\frac{b}{Ct}\right)^n, \text{ for } \theta < \frac{b}{CT} \tag{37}
$$

where *t* is the exit residence time representing the actual traveling time of the component of the fluid. The value of *b* for the combined chlorine was about 4 and 2.8 for total coliform and fecal coliform, respectively, and the value of *n* was about 3 for both total and fecal coliform [26]. *T* is the theoretical residence time (i.e., hydraulic retention time (HRT)), measured as the volume divided by the flow rate.

The overall survival fraction of bacteria can be determined by integrating the survival fraction of a series of batch reactors over the *E<sup>θ</sup>* function, if the reaction in a continuous flow reactor. In this case, the average disinfection efficiency (i.e., 1 – (N/No)) is expressed as follows [25]:

$$1 - \frac{N}{N\_0} = 1 - \int\_0^\sigma \left[\frac{N}{N\_0}\right]\_{\text{batch}} E\_\theta d\theta \tag{38}$$

Therefore, the average disinfection efficiency can be calculated by the given chlorine dose (*C*), dimensionless diffusion coefficient (*d*), disinfection coefficients (*b* and *n*), and residence time distribution function (*Eθ* and *dθ*) using Equation (39) [25]:

$$\begin{split} 1 - \int\_{0}^{\nu} \left[ \frac{N}{N\_{0}} \right]\_{\text{batch}} E\_{\theta} d\theta &= 1 - \int\_{0}^{b/CT} \frac{1}{\sqrt{4\pi\theta d}} \exp\left(\frac{-\left(1-\theta\right)^{2}}{4\theta d}\right) d\theta \\ &+ \int\_{b/CT}^{\nu} \left(\frac{b}{CT}\right)^{u} \frac{1}{\sqrt{4\pi\theta d}} \exp\left(\frac{-\left(1-\theta\right)^{2}}{4\theta d}\right) d\theta \end{split} \tag{39}$$

#### *4.1.3. Formation kinetics of THMs*

0.14 / *<sup>k</sup> <sup>d</sup>*

According to Collins and Selleck [26], the die-off of microorganisms in a batch reactor due to

q

q

1, for *N b N CT*

 , for *n Nb b N Ct CT*

where *t* is the exit residence time representing the actual traveling time of the component of the fluid. The value of *b* for the combined chlorine was about 4 and 2.8 for total coliform and fecal coliform, respectively, and the value of *n* was about 3 for both total and fecal coliform [26]. *T* is the theoretical residence time (i.e., hydraulic retention time (HRT)), measured as the

The overall survival fraction of bacteria can be determined by integrating the survival fraction of a series of batch reactors over the *E<sup>θ</sup>* function, if the reaction in a continuous flow reactor. In this case, the average disinfection efficiency (i.e., 1 – (N/No)) is expressed as follows [25]:

> q q

( )

q

q

2

<sup>2</sup> /

1 1

ò (38)

 q

(39)

( )

q

q

q

0 00 batch 1 1 *N N E d N N*

Therefore, the average disinfection efficiency can be calculated by the given chlorine dose (*C*), dimensionless diffusion coefficient (*d*), disinfection coefficients (*b* and *n*), and residence time

1 1

*<sup>b</sup> <sup>d</sup> CT d d*

*<sup>N</sup> E d <sup>d</sup> N d d*

pq

 é ù - =- ê ú ë û

¥

 é ù ê ú = <

where k is a nonideality coefficient of the reactor.

disinfection can be expressed by Equations (36) and (37):

0 batch

0 batch

distribution function (*Eθ* and *dθ*) using Equation (39) [25]:

0 0 batch 0

ò ò

q

q

<sup>1</sup> <sup>1</sup> exp <sup>4</sup> <sup>4</sup>

*n*

exp <sup>4</sup> <sup>4</sup>

æ ö æ ö - - <sup>+</sup> ç ÷ ç ÷ ç ÷ è ø è ø

pq

*b CT*

æ ö é ù - - - = - ç ÷ ê ú ç ÷ ë û è ø

/

¥

*b CT*

ò

¥

*4.1.2. Disinfection kinetics*

262 Mass Transfer - Advancement in Process Modelling

volume divided by the flow rate.

*L W* <sup>=</sup> (35)

ë û (36)

é ù æ ö ê ú = < ç ÷ ë û è ø (37)

The reaction chemistry between chlorine and aquatic humic precursors to form THMs can be expressed by the following short-hand way [25]:

$$\text{CL}\_2 + m^\text{"TOC} \text{(precursors)} \rightarrow \text{THMs} + \text{other products} \tag{40}$$

Based on the above process chemistry, the kinetics of chlorine decay and the THMs formation can be expressed by Equations (41) and (42), respectively:

$$-\frac{d\left[\text{Cl}\_2\right]}{dt} = k\_1 \left[\text{Cl}\_2\right] \left[\text{TOC}\right]^{n'}\tag{41}$$

$$\frac{d\left[\text{THMs}\right]}{dt} = k\_2 \left[\text{Cl}\_2\right] \left[\text{TOC}\right]^n \tag{42}$$

When reactions take place in a continuous flow reactor, the overall change of free chlorine and THMs can be obtained by integration of the batch concentration over the entire exit time distribution (E*θ*) as follows:

$$\begin{aligned} \left(\frac{\left[\text{Cl}\_2\right]}{\left[\text{Cl}\_2\right]\_0}\right) &= \int\_0^\varepsilon \left(\frac{\left[\text{Cl}\_2\right]}{\left[\text{Cl}\_2\right]\_0}\right)\_{\text{bach}} E\_\vartheta d\theta = \\ &= \int\_0^\varepsilon \frac{1}{\sqrt{4\pi\theta d}} \exp\left(\frac{-\left(1-\theta\right)^2}{4\theta d}\right) \exp\left(-k\_\text{i} T\theta \left[\text{TOC}\right]\right) d\theta \end{aligned} \tag{43}$$

$$\begin{aligned} \text{[THMs]} &= \mathop{\rm l}\limits\_{0}^{\text{w}} \text{[THMs]} \mathop{\rm l}\nolimits\_{\text{buth}} E\_{\vartheta} d\theta = \\ &= \mathop{\rm l}\limits\_{0}^{\text{w}} \mathop{\rm l}\nolimits \left[ \mathop{\rm TOC} \right]^{\text{w}} \left[ \mathop{\rm Cl}\nolimits\_{\text{l}} \right] \exp \left( - \mathop{\rm l}\nolimits \left[ \mathop{\rm TOC} \right] \right) \frac{1}{\sqrt{4\pi\theta d}} \exp \left( \frac{-\left(1-\theta\right)^{2}}{4\theta d} \right) d\theta \end{aligned} \tag{44}$$

#### **4.2. Engineering practice**

#### *4.2.1. Determination of coefficients in disinfection kinetic model*

In practice, the total coliform and total bacteria count can be measured by the membrane filter method and heterotrophic plate count method, respectively. The effect of chlorine doses and contact time on the survival of total coliform and total bacteria can be determined, as shown in Figures 6 (a) and (b), respectively.

**Figure 6.** Effect of applied chlorine and contact time on survival of (a) total coliform and (b) total count. □ Cl2, 1.2 mg/L; ■ Cl2, 2.0 mg/L; ΔCl2, 1.8 mg/L; ▲Cl2, 0.9 mg/L. (Adapted from Chiang et al. [25])

#### *4.2.2. Prediction of THMs Formation*

As shown in Figure 7, the effect of the contact time (i.e., 0–70 min) and the chlorine dose (i.e., 0.5–8.0 mg/L) on the THM formation was presented at a TOC of 3 mg/L with an L/W ratio of 20. The THM concentrations increase linearly with the contact time over the range of 0 to 70 min. By knowing the water quality parameters and experimental conditions, it is possible to predict the THM formation in the effluent of the chlorine contact chamber. To keep the THM formation below 100 µg/L at the water treatment plant, it is necessary to maintain the chlorine dose below 2 mg/L for a contact basin with an L/W ratio of 20.

**Figure 7.** Effect of applied chlorine dose ad contact time on THM formation at inlet TOC concentration of 3.0 mg/L and an L/W ratio of 20. (Modified from Chiang et al. [25])

#### *4.2.3. Balancing disinfection efficiency and THM formation*

**Figure 6.** Effect of applied chlorine and contact time on survival of (a) total coliform and (b) total count. □ Cl2, 1.2

As shown in Figure 7, the effect of the contact time (i.e., 0–70 min) and the chlorine dose (i.e., 0.5–8.0 mg/L) on the THM formation was presented at a TOC of 3 mg/L with an L/W ratio of 20. The THM concentrations increase linearly with the contact time over the range of 0 to 70 min. By knowing the water quality parameters and experimental conditions, it is possible to predict the THM formation in the effluent of the chlorine contact chamber. To keep the THM formation below 100 µg/L at the water treatment plant, it is necessary to maintain the chlorine

**Figure 7.** Effect of applied chlorine dose ad contact time on THM formation at inlet TOC concentration of 3.0 mg/L and

mg/L; ■ Cl2, 2.0 mg/L; ΔCl2, 1.8 mg/L; ▲Cl2, 0.9 mg/L. (Adapted from Chiang et al. [25])

dose below 2 mg/L for a contact basin with an L/W ratio of 20.

*4.2.2. Prediction of THMs Formation*

264 Mass Transfer - Advancement in Process Modelling

an L/W ratio of 20. (Modified from Chiang et al. [25])

As the aforementioned, balancing the disinfection efficiency (in Section 2.1) and THM formation (in Section 2.2) is an important task in chlorination process. Table 5 provides the engineering practice (exemplified by Case A wastewater treatment plant) to balance the disinfection efficiencies and THM formation, in terms of contact time, L/W ratio, and applied chlorine doses. It can be used to determine the applied chlorine dose needed to meet a specified degree of inactivation (e.g., 99.99 % kill) and level of THM formation. In general, the THM formation increases rapidly with the increases of contact time. Conversely, the disinfection efficiency of chlorine increases gradually as both contact time and chlorine dosage increase.

It is noted that the applied chlorine dose at about 1.0 mg/L exhibits the most economically feasible way to control the THM formation and to maintain the disinfection efficiency at the Case A water treatment plant. Under these conditions, the Case A plant is able to achieve an acceptable level of residual chlorine (i.e., 0.5 mg/L) at the THM formation below 100 µg/L.


**Table 5.** Disinfection efficiency and THM formation during chlorination at TOC = 3 mg/L in a real water treatment plant (Case A) (modified from Chiang et al. [25])

In addition, the THM concentration and disinfection efficiency associated with a given contact time, L/W ratio, and applied chlorine dose are also presented in Table 5. Prior to chlorine input to the solution, reducing the TOC concentration is definitely the best engineering practice (BEP) for controlling the THM formation, rather than improving the hydraulics through the geo‐ metric design of the reactor. On the other hand, the results imply that the geometry of a chlorine contactor can effectively control the disinfection efficiency (not on the formation of THMs), especially for large chlorine doses and a high L/W ratio.

Furthermore, in order to reduce the formation of THM, the ozonation process is sometimes utilized in disinfection for the purpose of eliminating bacteria, viruses, cysts, as well as organic compounds, especially, the precursors of disinfection DBPs [27]. The performance of the ozone disinfection processes is evaluated by the residual ozone and contact time. Therefore, factors such as the efficiency of gas–liquid contact, chemical reactivity of the raw water toward ozone, susceptibility of microorganisms, and hydrodynamic characteristics of the contactor can influence the disinfection performance of an ozonation process. Several researchers have shown that ozonation prior to chlorination can lower the formation potential of THM (triha‐ lomethanes) and HAA (haloacetic acid) [28, 29].

#### **5. Carbonation process**

#### **5.1. Model development**

#### *5.1.1. Process chemistry*

Carbonation is generally classified as a heterogeneous reaction containing gas, liquid, and solid phases. CO2 fixation by accelerated carbonation has been regarded as mass transfer limited reaction (i.e., film-diffusion controlled) according to the findings reported in the literature [30-32]. Therefore, different types of approaches such as physical intensification [33] and biological or chemical activation [34, 35] were recently carried out to improve the mass transfer and reaction kinetics. Based on this idea, a rotating packed bed (RPB) reactor has been introduced to improve the mass transfer rate among phases due to its high centrifugal forces and great micro-mixing ability. RPB, the so-called high-gravity (sometimes called the "HI‐ GEE") process, is able to generate high acceleration via centrifugal force so the mass transfer between gas and liquid and even between liquid and solid can be enhanced. Since RPB can provide a mean acceleration of hundreds, and even thousands, of times greater than the force of gravity, it can effectively lead to the formation of thin liquid films and micro- or nanodroplets [36-38]. Therefore, the volumetric gas–liquid mass transfer coefficients (KGa) are an order of magnitude higher than those in a conventional packed bed, leading to dramatic reductions in equipment size over that required for equivalent mass transfer in a gravity-flow packed bed [36, 38, 39].

According to the findings in the literature [40], the CO2 consumption in bulk solution was mainly attributed (over 98 % contribution) to carbonation reaction with calcium ions in solution leached from the steel slag to form calcium carbonate precipitates, where the stoi‐ chiometric formula was briefly presented as Equation (45). In this case, the stoichiometric coefficient (*b*) is assigned a value of one.

$$\text{CaO}\_{2(\text{fluid})} + \text{b }\text{CaO}\_{(\text{solid})} \rightarrow \text{CaCO}\_{3(\text{s})} \tag{45}$$

From the gas-phase point of view, the CO2 removal efficiency (η) for a carbonation process can be calculated as Equation (46):

$$\log\left(\%\right) = \left\lfloor 1 - \frac{\left(\rho\_{\text{CO}\_2,o}\,\text{Q}\_{\text{G},o}\,\text{C}\_{\text{G},o}\right)}{\rho\_{\text{CO}\_2,i}\,\text{Q}\_{\text{G},i}\,\text{C}\_{\text{G},i}}\right\rfloor \times \% \tag{46}$$

where the *ρCO*2,*<sup>i</sup>* and *ρCO*2,*<sup>o</sup>* are the densities of inflow and outflow streams of CO2, respectively. *QG*,*<sup>i</sup>* and *QG*,*<sup>o</sup>* are the flow rates of inflow and outflow streams, respectively. *CG*,*<sup>i</sup>* and *CG*,*<sup>o</sup>* are the CO2 concentration in inflow and outflow streams, respectively.

From the solid-phase point of view, the carbonation conversion (XB, %) of the solid can be defined as the ratio of reacted CaO fraction with CO2 to the original total available CaO content in solid before carbonation., as expressed in Equation (47):

$$X\_{B} \left(\%\right) = \left[\frac{\Delta m\_{\text{CO}\_{2}}}{MW\_{\text{CO}\_{2}}} \times \frac{MW\_{\text{Ca}}}{CaO\_{\text{total}}}\right] \times \% \tag{47}$$

where ΔmCO2 is the weight gain due to CO2 carbonation per dry weight of the solid sample. MWCaO and MWCO2 are the molecular weight of CaO and CO2, respectively. CaOtotal is the total available CaO content in the solid before carbonation.

#### *5.1.2. Rate-determining steps*

disinfection processes is evaluated by the residual ozone and contact time. Therefore, factors such as the efficiency of gas–liquid contact, chemical reactivity of the raw water toward ozone, susceptibility of microorganisms, and hydrodynamic characteristics of the contactor can influence the disinfection performance of an ozonation process. Several researchers have shown that ozonation prior to chlorination can lower the formation potential of THM (triha‐

Carbonation is generally classified as a heterogeneous reaction containing gas, liquid, and solid phases. CO2 fixation by accelerated carbonation has been regarded as mass transfer limited reaction (i.e., film-diffusion controlled) according to the findings reported in the literature [30-32]. Therefore, different types of approaches such as physical intensification [33] and biological or chemical activation [34, 35] were recently carried out to improve the mass transfer and reaction kinetics. Based on this idea, a rotating packed bed (RPB) reactor has been introduced to improve the mass transfer rate among phases due to its high centrifugal forces and great micro-mixing ability. RPB, the so-called high-gravity (sometimes called the "HI‐ GEE") process, is able to generate high acceleration via centrifugal force so the mass transfer between gas and liquid and even between liquid and solid can be enhanced. Since RPB can provide a mean acceleration of hundreds, and even thousands, of times greater than the force of gravity, it can effectively lead to the formation of thin liquid films and micro- or nanodroplets [36-38]. Therefore, the volumetric gas–liquid mass transfer coefficients (KGa) are an order of magnitude higher than those in a conventional packed bed, leading to dramatic reductions in equipment size over that required for equivalent mass transfer in a gravity-flow

According to the findings in the literature [40], the CO2 consumption in bulk solution was mainly attributed (over 98 % contribution) to carbonation reaction with calcium ions in solution leached from the steel slag to form calcium carbonate precipitates, where the stoi‐ chiometric formula was briefly presented as Equation (45). In this case, the stoichiometric

From the gas-phase point of view, the CO2 removal efficiency (η) for a carbonation process can

,, , ,, ,

*Q C Q C*

( ) ( ) <sup>2</sup> 2

r

r

 %1 % *CO o G o G o CO i G i G i*

é ù = ê ú - ´ ë û

2 fluid ( ) (solid) 3 s( ) CO + b CaO CaCO ® (45)

(46)

lomethanes) and HAA (haloacetic acid) [28, 29].

**5. Carbonation process**

266 Mass Transfer - Advancement in Process Modelling

**5.1. Model development**

*5.1.1. Process chemistry*

packed bed [36, 38, 39].

coefficient (*b*) is assigned a value of one.

h

be calculated as Equation (46):

The shrinking core model (SCM) has been utilized to determine the rate-limiting steps in a heterogeneous reaction because of its conceptual and mathematical simplicity [41]. In the SCM, it is assumed that the reaction occurs at the outer skin of solid particle and then proceeds at a narrow front [24]. The narrow front moves into the particle and leaves behind completely reacted product layer and/or reactive-species-depleted rims (i.e., ash layer). An unreacted core of material exists at any time, which shrinks in size during the reaction. Based on the above assumption, three factors might affect the reaction (i.e., rate-limiting steps) including (1) fluidfilm diffusion, (2) ash-layer (or product layer) diffusion, and (3) chemical reaction at the unreacted-core surface, as illustrated in Figure 8.

**Figure 8.** Schematic diagram of mass transfer among gas, liquid, and solid phases in the case of basic oxygen furnace slag (BOFS) carbonation.

When the reactant transfer through the boundary layer of the liquid–solid interface is the ratelimiting step of a reaction, the relationship between the conversion (XB, %) and reaction time (t, s) can be correlated with Equation (48):

$$t = \frac{\rho\_{\text{B}}R}{\Re b k\_{\text{e}} \mathbb{C}\_{A\_{\text{g}}}} X\_{\text{B}} = \tau \ X\_{\text{B}} \tag{48}$$

where *ρB* (mole cm-3) and *R* (cm) are the molar density and the radius of the particle, respec‐ tively, *CAg* (mole cm-3) is the CO2 concentration in the liquid phase, *ke* (mole m-2 Pa-1 s-1) is the mass transfer coefficient between the fluid and particle, and *τ* (s) is the time for complete conversion to a product particle.

When the reactant transfer through the ash layer is the rate-limiting step, the dependence of the reaction time on the conversion and the effective diffusivity of the reactant in the ash layer (De, cm2 s-1) can be described by Equation (49):

$$t = \frac{\rho\_{\rm B} \mathcal{R}^2}{6bD\_e \mathcal{C}\_{A\_{\rm g}}} \left[ 1 - \Im \left( 1 - X\_{\rm B} \right)^{2/3} + 2 \left( 1 - X\_{\rm B} \right) \right] = \tau \left[ 1 - \Im \left( 1 - X\_{\rm B} \right)^{2/3} + 2 \left( 1 - X\_{\rm B} \right) \right] \tag{49}$$

For simplicity, the particle radius (*R*) is generally assumed to remain constant in the literature. However, in reality, the *R* value is not a constant because the thickness of ash layer changes with the reaction, also affecting the diffusivity of the gaseous reactant in the ash layer. Sohn and Szekely [42] have introduced this phenomenon into SCM to make it more general to the fluid–solid reaction. In their recent report [43], a "Z" factor was incorporated with the governing equation of SCM for pore diffusion control, as shown in Equation (50). Z is defined as the volume of product solid formed from a unit volume of reactant solid, where both volumes include those of pores. In this case, the volume of the solid product is different from that of the solid reactant, which means R changes with time:

$$t = \frac{\rho\_{\rm B} \mathcal{R}^2}{2bD\_c \mathcal{C}\_{A\_{\rm g}}} \left[ \frac{Z - \left(Z - \left(Z - 1\right)\left(1 - X\_{\rm g}\right)\right)^{2/3}}{Z - 1} - \left(1 - X\_{\rm g}\right)^{2/3} \right] \tag{50}$$

If the *R* value is assumed to remain constant, the Z, herein, should be assigned a value of 1, i.e., the total volume of solid product is the same as that of solid reactant. Thus, Equation (50) will be simplified to Equation (49) by applying L'Hospital's rule.

When the chemical reaction between reactants is the rate-limiting step, the relationship between the conversion and reaction time can be determined by Equation (51):

$$t = \frac{\rho\_{\text{g}}R}{bk\,\text{\AA}} \left[ 1 - \left( 1 - X\_{\text{g}} \right)^{1/3} \right] = \,\,\,\tau \left[ 1 - \left( 1 - X\_{\text{g}} \right)^{1/3} \right] \tag{51}$$

where k" (s-1) is the first-order rate constant for the surface reaction.

#### *5.1.3. Mass transfer coefficients*

When the reactant transfer through the boundary layer of the liquid–solid interface is the ratelimiting step of a reaction, the relationship between the conversion (XB, %) and reaction time

*B B*

where *ρB* (mole cm-3) and *R* (cm) are the molar density and the radius of the particle, respec‐ tively, *CAg* (mole cm-3) is the CO2 concentration in the liquid phase, *ke* (mole m-2 Pa-1 s-1) is the mass transfer coefficient between the fluid and particle, and *τ* (s) is the time for complete

When the reactant transfer through the ash layer is the rate-limiting step, the dependence of the reaction time on the conversion and the effective diffusivity of the reactant in the ash layer

> ( ) ( ) ( ) ( ) <sup>2</sup> 2/3 2/3 1 31 21 1 31 21

For simplicity, the particle radius (*R*) is generally assumed to remain constant in the literature. However, in reality, the *R* value is not a constant because the thickness of ash layer changes with the reaction, also affecting the diffusivity of the gaseous reactant in the ash layer. Sohn and Szekely [42] have introduced this phenomenon into SCM to make it more general to the fluid–solid reaction. In their recent report [43], a "Z" factor was incorporated with the governing equation of SCM for pore diffusion control, as shown in Equation (50). Z is defined as the volume of product solid formed from a unit volume of reactant solid, where both volumes include those of pores. In this case, the volume of the solid product is different from

té ù é ù = -- +- = -- +- ê ú ë û ê ú ë û (49)

> ( ( ) ( )) ( ) 2/3 <sup>2</sup> 1 1 2/3

é ù --- - ê ú <sup>=</sup> - - - ë û

If the *R* value is assumed to remain constant, the Z, herein, should be assigned a value of 1, i.e., the total volume of solid product is the same as that of solid reactant. Thus, Equation (50)

When the chemical reaction between reactants is the rate-limiting step, the relationship

( ) ( ) 1/3 1/3 " 11 11 *<sup>B</sup>*

t

*B B*

é ùé ù = -- = -- ê úê ú ë ûë û (51)

1

*B*

(50)

*<sup>R</sup> t XX XX*

*BB BB*

t

(48)

 3 *B*

*e Ag <sup>R</sup> t XX bk C* r= =

(t, s) can be correlated with Equation (48):

268 Mass Transfer - Advancement in Process Modelling

conversion to a product particle.

*e Ag*

*bD C* r

s-1) can be described by Equation (49):

that of the solid reactant, which means R changes with time:

*e Ag*

r

2 1

will be simplified to Equation (49) by applying L'Hospital's rule.

*Ag*

where k" (s-1) is the first-order rate constant for the surface reaction.

*bk C* r

*bD C Z*

*B B*

between the conversion and reaction time can be determined by Equation (51):

*<sup>R</sup> tX X*

*R ZZZ X t X*

(De, cm2

6 *B* For determining the *KGa* value in a heterogeneous system containing the gas, liquid, and solid phases, the following assumptions are generally made for simplicity: (1) the effect of an inclined gas–liquid interface is neglected, (2) solid distribution throughout the bed is uniform, (3) the concentration of the liquid at the particle surface is equal to the saturation concentration of the solution (i.e., the mass transfer between the liquid and the solid is neglected), and (4) the changes in particle size and particle surface area are neglected. In this case, the *KGa* in a packed bed can be determined by the two-film theory as Equation (52):

$$\frac{1}{K\_{\rm c}a\_{\varepsilon}} = \frac{1}{k\_{\rm c}a\_{\varepsilon}} + \frac{H}{I(k\_{\rm c}a\_{\varepsilon})} \tag{52}$$

where kG is the gas-side mass transfer coefficient (m s-1), kL is the liquid-side mass transfer coefficient (m s-1), H is the Henry's law constant, and I is the enhancement factor.

On the other hand, the driving force between the saturated CO2 concentration in the bulk gas and the CO2 concentration on the surface of liquid film can be determined by shell mass balance over a thin film of fluid with the RPB as shown in Equations (53) and (54):

$$\frac{1}{\rho\_{\text{CO}\_2}} \frac{\text{d}M\_{\text{C}}}{\text{d}V} = \text{K}\_{\text{C}} \text{a}\_{\text{s}} \left( \text{C}\_{\text{G}}^{\text{\*}} - \text{C}\_{\text{G}}^{\text{\*}} \right) \tag{53}$$

$$\mathbf{d}V = 2\pi r \hbar \cdot \mathbf{d}r\tag{54}$$

where ρco2 is the CO2 mass density at the temperature of gas streams, MG is the gas mass flow rate (kg s-1), CG \* is the saturation concentration of CO2 in solution (mg/L), CG ' is the concentra‐ tion of CO2 in solution (mg/L), V is the volume of packed bed (m3 ), h is the height of packing bed (m), and r is the radius of packed bed (m).

In this case, by substitution of Equation (54) into Equation (53), the *KGa* can be determined as Equation (55):

$$K\_{\rm G}a\_{\varepsilon} = \frac{M\_{\rm G}}{\rho\_{\rm G}\hbar\pi\left(r\_{\rm o}^{2} - r\_{i}^{2}\right)} \left(\text{NTU}I\_{\rm G}\right) = \frac{Q\_{\rm G}}{\hbar\pi\left(r\_{\rm o}^{2} - r\_{i}^{2}\right)}\ln\left(\frac{\mathcal{C}\_{\rm G,i}}{\mathcal{C}\_{\rm G,o}}\right) \tag{55}$$

Furthermore, Kelleher and Fair (1996) have obtained an overall volumetric mass transfer coefficient in terms of the area of transfer unit (ATU) for the gas side from the literature [44]. The height of transfer unit (HTU) and ATU can be calculated from the experimental data using Equations (56) and (57), respectively [45]:

$$\text{HTU} = \frac{r\_o - r\_i}{\text{NTU}} = \frac{r\_o - r\_i}{\ln(\mathbb{C}\_{G,i} / \mathbb{C}\_{G,o})} \tag{56}$$

$$\text{ATU} = \frac{\pi \left(r\_o^2 - r\_i^2\right)}{\text{NTU}} = \frac{\pi \left(r\_o^2 - r\_i^2\right)}{\ln \left(\mathbb{C}\_{G,i} / \mathbb{C}\_{G,o}\right)}\tag{57}$$

It is noted that the *KGa* value in an RPB are an order of magnitude greater than those in a conventional packed bed. In other words, the HTU value of an RPB is smaller than that of conventional packed bed. Therefore, the volume of the RPB reactor for achieving a certain degree of performance is much smaller than that of a conventional reactor such as autoclave and slurry reactors [38, 39]. Several operating factors will affect the performance of RPB reactor. In general, KGa increases with the increases of the gas flow rate, the liquid flow rate, and mainly the rotor speed [46].

#### **5.2. Engineering practice**

#### *5.2.1. Determination of diffusion coefficients*

According to the findings reported by Changet al. [47], the accelerated carbonation of the basic oxygen furnace slag (BOFS) in an RPB is controlled by the ash-layer diffusion mechanism because the experimental data exhibit a good correlation with Equation (51). Similar findings were observed in the literature [48], which suggests that the carbonation reaction should be ash-layer diffusion controlled. The particles exhibited an ash layer of calcium carbonate (CaCO3) product after carbonation, which could be classified as a shrinking core. As shown in Figure 9, the *De* values of CO2 gas through the ash layer are in the range of 5.47×10-7 to 1.49×10-6 cm2 s-1, increasing with the increase of reaction temperature.

**Figure 9.** Variation of *D*e with reaction temperature for BOFS carbonation based on the SCM (operating conditions: PCO2, 1 bar; flow rate, 1.2 L min-1; Dp, ~ 62 µm; L/S, 20 mL/g-1) (adapted from Chang et al. [47])

In addition, Changet al. [41]) reported an effective diffusivity of steelmaking slag in a slurry reactor of 2.9×10-7 to 7.3×10-7 cm2 s-1 according to the results of the SCM. It was found that the rate of CO2 molecular diffusion was not significantly enhanced by the RPB because the very similar temperatures were operated in the two reactors. Therefore, the *De* values measured in a slurry reactor exhibit the same magnitude of 10-6 cm2 s-1 to that in an RPB. Equation (58) shows the relationship between *De* value (L2 T-1) and *KGa* (L T-1) value:

$$K\_G = \frac{D\_\kappa}{\delta} \tag{58}$$

where *δ* (L) is the thickness of fluid film. Only a slight discrepancy in *De* values between an RPB and a slurry reactor was found; however, the RPB reactor can generate a thinner liquid film (*δ*), thereby resulting in a higher mass transfer rate.

#### *5.2.2. Determination of overall gas-phase mass transfer coefficients*

It is noted that the *KGa* value in an RPB are an order of magnitude greater than those in a conventional packed bed. In other words, the HTU value of an RPB is smaller than that of conventional packed bed. Therefore, the volume of the RPB reactor for achieving a certain degree of performance is much smaller than that of a conventional reactor such as autoclave and slurry reactors [38, 39]. Several operating factors will affect the performance of RPB reactor. In general, KGa increases with the increases of the gas flow rate, the liquid flow rate, and mainly

According to the findings reported by Changet al. [47], the accelerated carbonation of the basic oxygen furnace slag (BOFS) in an RPB is controlled by the ash-layer diffusion mechanism because the experimental data exhibit a good correlation with Equation (51). Similar findings were observed in the literature [48], which suggests that the carbonation reaction should be ash-layer diffusion controlled. The particles exhibited an ash layer of calcium carbonate (CaCO3) product after carbonation, which could be classified as a shrinking core. As shown in Figure 9, the *De* values of CO2 gas through the ash layer are in the range of 5.47×10-7 to

**Figure 9.** Variation of *D*e with reaction temperature for BOFS carbonation based on the SCM (operating conditions:

PCO2, 1 bar; flow rate, 1.2 L min-1; Dp, ~ 62 µm; L/S, 20 mL/g-1) (adapted from Chang et al. [47])

s-1, increasing with the increase of reaction temperature.

the rotor speed [46].

1.49×10-6 cm2

**5.2. Engineering practice**

*5.2.1. Determination of diffusion coefficients*

270 Mass Transfer - Advancement in Process Modelling

Based on the above illustration, the KGa value in an RPB can be estimated by the key operating parameters including gas diffusivity, gas superficial velocity, gas density, gas viscosity, liquid superficial velocity, liquid density, liquid viscosity, centrifugal acceleration, total specific surface area of packings, and effective diameter of packings. Therefore, the mass transfer characteristics of RPB (such as KGa value and HTU) can be obtained by the nonlinear regression of experimental data [49]:

$$K\_{\rm G}a\_{\rm \epsilon} = 0.01 \left(\frac{a\_{\rm \epsilon} D\_{\rm G}}{d\_{\rm \epsilon}}\right) Re\_{\rm G} \,^{-1.16} \,\mathrm{Gr}\_{\rm G} \,^{0.33} \,\mathrm{Re}\_{\rm L} \,^{2.12} \tag{59}$$

$$HTLI = 0.0003 \left( \frac{a\_t D\_G}{d\_p} \right) Re\_G^{-2.16} \text{ } Gr\_G^{-0.33} \text{ } Re\_L^{-2.12} \tag{60}$$

where the ranges of the dimensionless groups in this correlation should be as follows:

$$17.8 < Re\_{\odot} < 15.9\tag{61}$$

$$1.3 < \text{Re}\_L < 2.2\tag{62}$$

$$2.3 < Gr\_{\odot} < 26.8 \tag{63}$$

Table 6 presents the *KGa* value for high-gravity carbonation of BOFS-CRW, respectively. The rotation speed varied from 150 to 550 rpm, offering a centrifugal acceleration variation from 60 m/s2 to 770 m/s2 . The *KGa* values moderately increase with an increase of rotation speed (i.e., up to 300–500 rpm), indicating that the mass transfer resistance was reduced by an increasing rotation speed within this range. However, a reduction in *KGa* was observed if the rotation speed further increased beyond this range. This might be attributed to the fact that the extent of reduction in mass transfer resistances at higher rotation speed was compensated for by a reduction of the retention time, which was unfavorable to reaction.


**Table 6.** Candidates of optimal KGa solutions under different operating conditions and verified with theoretical model

On the other hand, several studies have utilized different kinetic models to determine the ratelimiting step of mineral carbonation [31, 41]. In fact, solid–liquid mass transfer is particularly important in mineral carbonation and, in many cases, the rate limiting factor [50, 51] because minerals in solid matrix dissolve partly and passive layers are formed, gradually increasing resistance to mass transfer and eventually leading to incomplete conversion. In general, the solid–liquid mass transfer coefficient is occasionally correlated as itself, where such correla‐ tions are specific to the system under consideration and are not generally applicable [50].

In spite of the significant differences between RPB and traditional packed column, penetration theory was still capable of describing the liquid-side mass transfer behavior fairly well in RPB [52]. These correlations are most often expressed in terms of dimensionless numbers in the form of a power series. For instance, penetration theory can be applied to RPB to yield as Equation (64):

$$k\_L = 0.92 \left(\frac{D\_L}{d\_p}\right) \text{Sc}\_L^{1/2} \text{ Re}\_L^{1/3} \text{Gr}\_L^{1/6} \tag{64}$$

where the Grashof number which represents the ratio of gravitational to viscous forces can be determined by Equation (65) and the *g* value can be replaced by the centrifugal acceleration term as shown in Equation (66):

Application of Mass Transfer Models in Environmental Engineering http://dx.doi.org/10.5772/60917 273

$$Gr\_L = gd\_p^3 \left(\frac{\rho\_L}{\mu\_L}\right)^2\tag{65}$$

$$\mathbf{g} = \mathbf{a}\_m = \left(\frac{r\_o^2 + r\_i^2}{2}\right)^{1/2} o r^2 \tag{66}$$

The Grashof (*Gr*) number is generally determined by the mean radius of the packed bed. Another commonly used prediction for liquid-side (kL) and gas-side (kG) mass transfer coefficient in a conventional packed column is that of Onda et al. (1968), as shown in Equations (67) and (68), respectively. Onda et al. suggested the constant 5.23 of Equation (68) should be best correlated by changing the constant into 2.00 for smaller packings (i.e., diameter is less than 1.5 cm) since the KGa data for packings smaller than 1.5 cm tend to decrease in the literatures [53].

$$k\_L = 0.0051 \left(\frac{a\_t}{a\_w}\right)^{2/3} Re\_L^{2/3} Sc\_L^{-1/2} \left(\frac{\rho\_L}{\mu\_L g}\right)^{-1/3} \left(a\_p d\_p\right)^{0.4} \tag{67}$$

and

60 m/s2

**No ω'a**

**a**

Equation (64):

term as shown in Equation (66):

to 770 m/s2

272 Mass Transfer - Advancement in Process Modelling

**(rpm)**

ω', rotation speed (rpm); QG, gas flow rate (m3

. The *KGa* values moderately increase with an increase of rotation speed (i.e.,

**G/L ratio (-) L/S ratio**

**(mL/g)**

**KGa value (s-1)**

up to 300–500 rpm), indicating that the mass transfer resistance was reduced by an increasing rotation speed within this range. However, a reduction in *KGa* was observed if the rotation speed further increased beyond this range. This might be attributed to the fact that the extent of reduction in mass transfer resistances at higher rotation speed was compensated for by a

reduction of the retention time, which was unfavorable to reaction.

**/min) Qsl <sup>a</sup>**

**(L/min)**

 462.1 0.33 9.33 35.4 10.12 0.672 462.7 0.34 9.33 36.4 10.61 0.650 505.1 0.82 6.50 126.2 12.35 0.715 487.3 0.44 9.17 48.0 11.21 0.680 534.4 0.72 6.33 113.7 12.77 0.632 360.2 0.42 9.00 46.7 10.06 0.699

/min); Qsl, slurry flow rate (L/min)

**Table 6.** Candidates of optimal KGa solutions under different operating conditions and verified with theoretical model

On the other hand, several studies have utilized different kinetic models to determine the ratelimiting step of mineral carbonation [31, 41]. In fact, solid–liquid mass transfer is particularly important in mineral carbonation and, in many cases, the rate limiting factor [50, 51] because minerals in solid matrix dissolve partly and passive layers are formed, gradually increasing resistance to mass transfer and eventually leading to incomplete conversion. In general, the solid–liquid mass transfer coefficient is occasionally correlated as itself, where such correla‐ tions are specific to the system under consideration and are not generally applicable [50].

In spite of the significant differences between RPB and traditional packed column, penetration theory was still capable of describing the liquid-side mass transfer behavior fairly well in RPB [52]. These correlations are most often expressed in terms of dimensionless numbers in the form of a power series. For instance, penetration theory can be applied to RPB to yield as

> 1/2 1/3 1/6 0.92 *<sup>L</sup> L L LL p <sup>D</sup> <sup>k</sup> Sc Re Gr d* æ ö ç ÷ ç ÷ è ø

where the Grashof number which represents the ratio of gravitational to viscous forces can be determined by Equation (65) and the *g* value can be replaced by the centrifugal acceleration

= (64)

**QG <sup>a</sup> (m3**

$$k\_G = 5.23 \left( a\_p D\_G \right) \operatorname{Re}\_G^{0.7} \operatorname{Sc}\_G^{1/3} \left( a\_p d\_p \right)^{-2} \tag{68}$$

Itisusuallydifficulttoobtainmass transfer coefficients separatedfromvolumetricmass transfer coefficients kLae and kGae since the effective interfacial area between the liquid and vapor phase isusuallynotknown[38].Several correlationshavebeenreportedtoestimate thewettedsurface area (aw), among which Tung and Mah [52] found that the aw/at predicted value by Perry and Chilton [54] is "reliable" under high-gravity RPB process as shown in Equation (69).

$$\frac{a\_w}{a\_t} = \frac{a\_c}{a\_t} = 1 - \exp[-1.45 \left(\frac{\sigma\_c}{\sigma\_L}\right)^{0.75} Re\_L^{0.1} \text{ } We\_L^{0.2} \text{ } Fr\_L^{-0.05} \text{ } \tag{69}$$

The assumption of the above equality is that for packing materials with small static holdup (i.e., large packing size), the wetted area (aw) may equal the interfacial area (ae). In addition, Basic (1992) has correlated the wetted surface area of packing by combining experimental data of liquid holdup and estimation of mean liquid film thickness, as shown in Equation (70):

$$\frac{d\sigma\_w}{d\sigma\_t} = 584 \text{ Re}\_L^{-1.03} \text{ } \text{We}\_L^{0.576} \text{ } Fr\_L^{0.123} \text{ } \tag{70}$$

#### *5.2.3. Balancing mass transfer performance and energy consumption*

Balancing the gas-phase mass transfer rate for the high-gravity carbonation process with a relativelylowerenergyconsumptionisalsoanimportanttaskforcarbonationprocess.Asshown in Figure 10, the favorable operating conditions can be determined via graphical presenta‐ tion, where the energy consumptions including rotation, blowers, air compressors, and pumps were measured during operation in terms of kWh per ton CO2 capture by the RPB. It suggests that a centrifugal acceleration should be maintained at 475 m/s2 for a relatively lower energy consumption (L1 → L2) and higher *KGa* value (L3 → L4). Moreover, the favorable G/L ratio should range between 40 and 55 for high-gravity carbonation process (by both L5 → L6 and R1 → R2 → R3). A further increase in G/L ratio up to 80 will lead to a low mass transfer perform‐ ance (i.e., KGa value) with high energy consumption for rotation and pumps, thereby exhibit‐ ing a poor CO2 removal efficiency and capacity.

**Figure 10.** Graphical presentation for determining the optimal *KGa* value with favorable centrifugal acceleration (i.e., rotation speed) and G/L ratio for high-gravity carbonation process (as indicated by red line) (courtesy of Pan et al. [49])

### **6. Summary**

*5.2.3. Balancing mass transfer performance and energy consumption*

that a centrifugal acceleration should be maintained at 475 m/s2

ing a poor CO2 removal efficiency and capacity.

274 Mass Transfer - Advancement in Process Modelling

Balancing the gas-phase mass transfer rate for the high-gravity carbonation process with a relativelylowerenergyconsumptionisalsoanimportanttaskforcarbonationprocess.Asshown in Figure 10, the favorable operating conditions can be determined via graphical presenta‐ tion, where the energy consumptions including rotation, blowers, air compressors, and pumps were measured during operation in terms of kWh per ton CO2 capture by the RPB. It suggests

consumption (L1 → L2) and higher *KGa* value (L3 → L4). Moreover, the favorable G/L ratio should range between 40 and 55 for high-gravity carbonation process (by both L5 → L6 and R1 → R2 → R3). A further increase in G/L ratio up to 80 will lead to a low mass transfer perform‐ ance (i.e., KGa value) with high energy consumption for rotation and pumps, thereby exhibit‐

**Figure 10.** Graphical presentation for determining the optimal *KGa* value with favorable centrifugal acceleration (i.e., rotation speed) and G/L ratio for high-gravity carbonation process (as indicated by red line) (courtesy of Pan et al. [49])

for a relatively lower energy

A clear understanding in process chemistry, reaction kinetics, and mass transfer is the essential requirement to achieve the best achievable technology. Many industrial and/or waste treat‐ ment processes are involved in diffusion-controlled mechanism. To increase process efficiency and reaction rate, the mass transfer of those processes should be improved, which can be approached by development of prediction model using theoretical theory such as Fick's law or penetration theory. For those heterogeneous processes, especially the system containing gas, liquid, and solid phase, assumptions should be appropriately made to simplify the complex governing equation. Lastly, from the viewpoint of engineering practice, the energy consumption of a process is also an important factor to optimize the overall process, thereby reducing the operation costs.

### **Author details**

Pen-Chi Chiang\* and Shu-Yuan Pan

\*Address all correspondence to: pcchiang@ntu.edu.tw

Graduate Institute of Environmental Engineering, National Taiwan University, Taiwan

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