Cwo of Phenol in an Trickle

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Water Research 37 (2003) 319–328

Catalytic wet oxidation of phenol in a trickle bed reactor over a

Pt/TiO2 catalyst

Clayton B. Maugans, Aydin Akgerman*

Chemical Engineering Department, Texas A&M University, College Station, TX 77843-3122, USA

Received 1 January 2002; received in revised form 1 June 2002; accepted 14 June 2002

Abstract

Catalytic wet oxidation of phenol was studied in a batch and a trickle bed reactor using 4.45% Pt/TiO2 catalyst in the

temperature range 150–2051C. Kinetic data were obtained from batch reactor studies and used to model the reaction

kinetics for phenol disappearance and for total organic carbon disappearance. Trickle bed experiments were then

performed to generate data from a heterogeneous flow reactor. Catalyst deactivation was observed in the trickle bed

reactor, although the exact cause was not determined. Deactivation was observed to linearly increase with the

cumulative amount of phenol that had passed over the catalyst bed. Trickle bed reactor modeling was performed using

a three-phase heterogeneous model. Model parameters were determined from literature correlations, batch derived

kinetic data, and trickle bed derived catalyst deactivation data. The model equations were solved using orthogonal

collocations on finite elements. Trickle bed performance was successfully predicted using the batch derived kinetic

model and the three-phase reactor model. Thus, using the kinetics determined from limited data in the batch mode, it is

possible to predict continuous flow multiphase reactor performance.

r 2002 Elsevier Science Ltd. All rights reserved.

Keywords: Phenol; Trickle bed reactors; Wet oxidation; Catalytic wet oxidation; Multiphase reactors

1. Introduction

Organic compounds are common pollutants in

industrial waste streams and waste sites. To remove

the pollutants there are many remediation technologies,

of which only reactive destruction results in the

mineralization of the waste. Bioremediation and incin-eration are two typical examples of reactive destruction

technologies. However, both these remediation technol-

ogies have their limitations and are not always the

optimal solution. Bioremediation is ideally suited for

low concentrations, with maximum concentrations as

low as 50 ppm in some examples [1]. Incineration is ideal

for highly concentrated liquid streams, typically around

350,000ppm or higher chemical oxygen demand (COD)

content before the energy requirements become self-

sustaining [2]. Neither process is optimal when the toxic

organic compound concentration falls between the

extremes. Some alternative destruction technologies

have also been studied in the field including wet air

oxidation (WAO), and supercritical water oxidation

(SCWO). WAO involves treating the aqueous organicswith air at elevated temperature and pressures [3–12]

and SCWO is the same, but occurring above the critical

point of water [13–20].

A natural extension to WAO is catalytic wet oxidation

(CWO), which is similar to WAO but with the addition

of a catalyst. The catalyst allows milder operating

conditions than WAO while yielding similar if not

superior kinetic performance. CWO research has been

conducted on a variety of organic compounds using

numerous catalysts with varying results [21–39]. CWO is

in use today as a treatment technology in certain

applications such as coal gas effluent oxidation in Japan

*Corresponding author. Tel.: +1-979-845-3375; fax: +1-

979-845-6446.

E-mail address:[email protected] (A. Akgerman).

0043-1354/02/$ - see front matter r 2002 Elsevier Science Ltd. All rights reserved.

PII: S 0 0 4 3 - 1 3 5 4 ( 0 2 ) 0 0 2 8 9 - 0



and Britain, as well as for cyanide wastewater oxidation

in Japan [27], and for destroying organic undesirables

in an ammonium sulfate crystal production facility in

Italy [9].

In this research heterogeneous CWO was studied as

an organic wastewater treatment technology using

phenol as a model compound. Phenol was selected

because it is a common industrial pollutant, is toxic, and

is difficult to biodegrade at high-concentration levels.

This model compound was studied in two different

reactor configurations; first a batch reactor was used to

determine technology viability as well as determination

of the reaction kinetics, then a flow reactor system was

developed primarily for model confirmation and scale-

up, but also to evaluate this technology in a industrially

more practical flow reactor system. One of the objectives

of this study was to determine whether trickle bedreactor operation can be predicted from kinetics in terms

of lumped parameters, such as total organic carbon

(TOC), determined in a batch reactor by a limited

number of experiments.

2. Reaction kinetics from batch experiments: phenol

destruction

Batch experiments originally performed using pow-

dered catalyst yielded the kinetic expression reported in

Maugans and Akgerman [30]. A reevaluation of

those data in this paper suggests that Eq. (1) is a more

suitable kinetic model. The rate expression was reder-

ived using the following mechanism based on reaction

between adsorbed phenol and disassociatively adsorbed

oxygen.

I: Ph þ SÃ"PhSÃ;

II: O2 þ 2SÃ"2OSÃ;

III: PhSÃ þ OSÃ-ASÃ þ SÃ;

IV: ASÃ"A þ SÃ:

If it is assumed that the surface reaction, reaction III is

the rate determining step, all other reactions are in

dynamic equilibrium, and rate IV is very fast, then the

following rate expression is obtained:

ratePh;intrinsic ¼qC Ph

W

V L

qt

¼ÀK 1C Ph

ffiffiffiffiffiffiffiffiffiffiC Oxy

p ð1 þ K PhC Ph þ

ffiffiffiffiffiffiffiffiffiffiffiffiffi ffiffiffiffiffiffiffiK OxyC Oxy

p Þ2

: ð1Þ

Parameters were determined from the original

batch reactor data reported by Maugans and Akgerman

[30–31] and are given in Table 1. Fig. 1 shows prediction

of experimental phenol conversion in the batch reactor

with Eq. (1).

Nomenclature

Ar Arrhenius preexponential factor

C i concentration of i (mol/L or ppm)

Deff effective diffusivity (m2/s)

E a activation energy (J/mol)hG gas holdup

hL liquid holdup

H Henry’s constant (atm/(mol solute/mol sol-

vent))

k reaction rate constant (units vary)

k La gas–liquid interfacial mass transfer coefficient

(1/min)

k LS liquid–solid interfacial mass transfer coeffi-

cient (L/min)

K reaction adsorption constant (L/mol)

P pressure (atm)

r radius (m)

R catalyst radius (m)

rate reaction rate (mol/min/g-cat)

RTP room temperature and pressure

t time (min)

T temperature (K)

vt total volumetric flowrate

W catalyst mass (g-cat)

Subscripts

A reactive TOC organic compounds

B unreactive TOC organic compounds

bed bedG gas

L liquid

Oxy oxygen

Ph phenol

S solid

TOC total organic carbon

Superscripts

0 initial0 dimensioness variable

* catalyst site

Greek Letters

e porosity

r density (g/L)

x catalyst deactivation

r 2 second-order gradient

C.B. Maugans, A. Akgerman / Water Research 37 (2003) 319–328320



3. Reaction kinetics from batch experiments: TOC

destruction

The overall TOC reduction mechanism was described

in terms of a parallel reaction scheme, which reduces to

the reactions below [30].

A -rateY

B

krateX

CO2:

In this reaction scheme, A denotes phenol and all other

intermediates produced from the oxidation of phenol

that further oxidize into refractory components B

(certain organic acids) or completely mineralize into

CO2. The initial value of A is known, since it is the same

as the initial TOC value. The initial value of B is

assumed to be 0 at time=0. The TOC reaction scheme

reduces to Eqs. (2) and (3) as shown by Maugans [31].

rateTOC;intrinsic ¼qC TOC

W

V L

qt

¼ À

K X

ð1 þ ffiffiffiffiffiffiffiffiffiffiffiffi ffiffiffiffiffiffiffiffi

K OxyC Oxy

p Þ2

!C A


p ;

ð2Þ

where C A concentration of phenol and all other

intermediates that further oxidize, is determined by

rateA;intrinsic ¼qC AW

V L

qt

¼ ÀK X

ð1 þ ffiffiffiffiffiffiffiffiffiffiffiffiffi ffiffiffiffiffiffiffi

K OxyC Oxy

p Þ2þ K Y

!C A


p ð3Þ

and the initial conditions C 0

A (initial phenol concentra-tion) and C TOC [0] are known (and equal). Kinetic

parameters for Eqs. (1)–(3) are also given in Table 1 and

Fig. 2 shows the prediction of the measured TOC

profiles in the batch reactor with Eqs. (2) and (3).

4. Trickle bed reactor studies

4.1. Materials

Phenol solution (90 wt% phenol 10 wt% water) was

obtained from Fisher Scientific. The catalyst was

powdered 4.45% Pt/TiO2 with a maximum diameter of

less than 105 mm, obtained from Engelhard. The binder

used with the catalyst, as described below, was a 40%

SiO2 solution of Ludoxt Binder. The oxygen source in

Table 1

Apparent Arrhenius constants for Eqs. (2)–(4)

Ari E ai (J/mol)

K 1 8.52Â 1012 (L1.5 g-cat/min/mol0.5) 110,000 (722,000)

(73.0Â 1012)

K Ph 5.76Â 1012 (L/mol) 100,000 (726,000)

(74.2Â 1011)

K Oxy 3.09Â 108 (L/mol) 67,000 (716,000)

(71.5Â 108)

K X 2.94Â 1014 (L1.5 g-cat/min/mol0.5) 125,000 (72300)

(71.4Â 1014)

K Y 6.42Â 107 (L1.5 g-cat/min/mol0.5) 79,000 (72400)

(73.0Â 107)

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 20 40 60 80 100 120

Time (min)

1 - C o n v e r s i o n

150 C

165 C

175 C

185 C200 C

Fig. 1. Batch reactor data, phenol conversion predictions with the kinetic model at different temperatures (P ¼ 48atm, W ¼ 2 g-cat/L,

C

0

ph ¼ 1150 ppm).

C.B. Maugans, A. Akgerman / Water Research 37 (2003) 319–328 321



these experiments was compressed air provided by

Brazos Valley Welding. All water used in experimenta-

tion and cleaning was distilled and deionized water.

TOC analysis reagents included J.T. Baker sodium

persulfate and EM 98% phosphoric acid. HPLC

analyses employed EM HPLC grade methanol and a

15 cm Supelcosil LC-8 column from Supelco. All glass

beads used in the reactor were Pyrext.

4.2. Equipment

Experiments were conducted in a trickle bed flow

reactor schematically represented in Fig. 3. All tubing

and fittings used were stainless steel. The gas flow rate

was adjusted and controlled by a mass flow controller.

The main pump (for water) was a variable speed,

reciprocating pump while the metering pump (for

phenol solution) was a smaller variable speed, recipro-

cating pump (Minipump). Heating was provided by

electric heaters located external to the reactor, which

heated a bed of aluminum beads in which the reactor

was immersed. Reactor temperature was monitored by asliding thermocouple in a thermowell. Two diaphragm-

type backpressure regulators (BPR) were used in series

to control reactor pressure. System pressure was

monitored with a Bordon tube pressure gauge. The

reactor effluent was cooled before expansion through

the BPRs using tap water and a double pipe heat

exchanger.

4.3. Experimental procedure

Deionized water was pumped at a constant flow rate

via the main pump. Into this flow stream, a stream of

concentrated phenol was added through the MiniPump.

Both pumps were reciprocating piston pumps, so to

dampen out pulses the feed stream was sent through a 1/

1600 tubing coil. The feed was then preheated in a 1/800

preheater coil immersed in the hot aluminum bed

surrounding the reactor. Air was used as the oxygen

source and flowed from the pressurized air tank,

through the mass flow controller, through the preheater,

and into the top of the reactor where it was mixed with

the phenol solution feed before entering the reactor. The

reactor configuration was such that the feed passed over

a bed of inert glass beads before entering the catalyst

zone. After exiting the reactor the effluent was cooled

and depressurized.

The reactor was loaded with 2.29 g of pelletized

catalyst and run continuously until the completion of

the flow reactor studies. During extended periods when

no data were taken, water was run through the reactor to

avoid disruptive cooling/heating and reactor shutdown.

Liquid flowrate was maintained constant at 12 g/min

and the air inlet flowrate was maintained at 171 mL/min

(measured at room temperature and pressure). Theflowrates were kept constant throughout the trickle bed

studies in order to maintain consistent hydrodynamics in

the reactor to eliminate the effect of reactor hydro-

dynamics on conversion. Space time was controlled by

adjusting the concentration of the phenol feed. Pressure

was maintained at approximately 4176 atm. The

temperature range studied was from 1501C to 2051C

and the feed concentration range varied from 200 to

1200ppm phenol. Because it was found that the

temperature profile along the catalyst bed of the reactor

was constant (711C), the thermocouple was left in place

at the bottom of the catalyst bed for all experiments.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 20 40 60 80 100 120

Time (min)

1 - C o n v e r s i o n

150 C

165 C

175 C

185 C

200 C

Fig. 2. Batch reactor data, TOC conversion predictions with the kinetic model at different temperatures (P ¼ 48 atm, W ¼ 2 g-cat/L.

C 0ph ¼ 1150 ppm).




4.4. Catalyst preparation

The research catalyst was supplied in a fine powder

form and unsuitable for fixed bed reactor experiments.

To pelletize the powder, 5.5 g of catalyst powder was

mixed with 6 mL of Ludoxt binder. The resulting paste

was then extruded, dried, crushed, sieved, and calcined

overnight at 3201C. This relatively low calcination

temperature was chosen since significant catalyst deac-

tivation was observed when calcining at temperatures

near 5001C.

The pellets were then tested in high temperature water

for physical stability and structural integrity. Below the

boiling point of water the catalyst was observed to be

unaffected by the hot water. At boiling, the catalyst

quickly disintegrated into a powdery form, due to the

formation and flashing of water vapor inside the pellet

pores. Table 2 gives the physical properties of the catalyst.In addition to Table 2, the following physical

properties were measured: median catalyst particle size

of 450 mm, bed density of 0.79 g/mL, bed porosity of

0.68, and solid material density of 2.4 g/(mL solid

mater). The pelletized catalyst contained 0.64 g of

4.45% Pt/TiO2 powdered catalyst per gram of pellets

(the remainder being binder).

4.5. Homogeneous reactions

In the flow reactor system the reactants were in a high

temperature reaction zone for a brief period of time

before and after the catalyst bed. This had the potential

of introducing error to the kinetic study as there was the

possibility of significant homogeneous wet oxidation

occurring outside the catalyst bed. Experiments were

therefore performed in which the catalyst bed was

replaced with similar sized inert glass beads. The reactor

was operated at 1851C, 41 atm, 12 g feed/min, 172 mL

(RTP)/min air, and a feed concentration of 417 ppm

phenol corresponding to 319 ppm TOC. Although

there was about 10% phenol conversion to certain

Fig. 3. Trickle bed reactor system.

Table 2

Catalyst surface areas

Stock

powder

LUDOX bound pellets

1101C dried Calcined

Single point surface

area P /P 0 (m2/g)

52.7 50.4

BET surface area

(m2/g)

51.7 56.1 53.7

BJH adsorption

(m2/g)

63.9 62.2

MJH desorption

(m2/g)

71.0

Cumulative pore V

(m3/g)

0.19 0.18

Avg. pore diameter

( (A)

116 118




intermediates, there was no reduction in TOC concen-

tration (Fig. 4) indicating that homogeneous oxidation

was not significant at the reaction conditions. The

slight reduction from the theoretical value of 319 ppmTOC to measured value of 300 ppm TOC in Fig. 4 is

probably due phenol conversion into intermediates,

which have less carbon mass fraction compared to

phenol.

5. Results: catalyst deactivation

Catalyst deactivation was observed in this study,

although the exact cause is uncertain. Fig. 5 shows

reactor performance at the reference condition through-

out the course of a continuous reactor study. Thereference condition was 1751C, 44 atm pressure and

460ppm initial phenol condition, the reaction was

repeated at this condition periodically. Fig. 5 shows a

linear decrease in conversion with respect to the

cumulative amount of phenol passed over the catalyst

bed. A linear regression of the above deactivation data

was performed and used in conjunction with the rate

expressions in the trickle bed model to account for

catalyst deactivation as shown in Eq. (4).

ratei ¼ 0:64ð1 À xÞratei ;intrinsic; ð4Þ

where the intrinsic rate expressions are those given by

Eqs. (1)–(3). x was found to be 0.0055Â cumulative

grams of phenol passed through the reactor. The factor

0.64 was included to account for the catalyst dilution

due to the binder. Therefore:

ratePh ¼ 0:64 ð1 À xÞ ratePh;intrinsic; ð5Þ

rateOxy ¼ 0:64 ð1 À xÞ rateoxy;intrinsic ¼ 6ratePh: ð6Þ

Eq. (6) was based on the assumption that 6 mol of

oxygen were consumed per mole of phenol. This was

based on TOC reduction rather than complete miner-

alization.

6. Mathematical modeling: phenol equations

A trickle bed reactor is a three-phase reactor, and it

was assumed that the reaction was only occurring on thesolid phase surface. It was necessary for the reactants to

transfer into the solid catalytic phase for reaction to

occur. To model this transfer a mass balance for each of

these phases was performed. Derivations can be found

in [31].

Gas phase:

hGrbedvt;GP 0Oxy

R1T W total

dP 0Oxy

dW 0

þ k LaP 0OxyP 0Oxy

H Oxyrmolar;water

À C 0Oxy;LC 0Oxy

!¼ 0; ð7Þ

liquid phase:

hLvt;LrbedC0Oxy

W total

dC 0Oxy

dW 0

À k LaP 0OxyP 0Oxy

H Oxyrmolar;water

À C 0Oxy;LC 0Oxy

!

þ3 fDeff ;Oxyð1 À eSÞC 0Oxy

R2

dC 0Oxy

dr0

r¼R

¼ 0; ð8Þ

hLvt;LrbedC 0Ph

W total

dC 0Ph

dW 0

þ3 fDeff ;Phð1 À eSÞC 0Ph

R2

dC 0Ph

dr0

r¼R

¼ 0; ð9Þ

solid phase:

Deff C 0Ph

Rr2C 0ph À rS ratePh ¼ 0; ð10Þ

Deff C 0Oxy

Rr2C 0Oxy À rS rateOxy ¼ 0: ð11Þ

Experimentally no temperature gradient along the

length of the reactor was observed. An internal energy

balance on the pellets showed no significant temperature

00.1

0.20.30.40.50.60.70.80.9

1

0 50 100 150

g Phenol Passed

C o n v e r s i o n

Phenol

TOC

Linear (TOC)

Fig. 5. Catalyst deactivation at base condition (T ¼ 1751C,

P ¼ 44 atm, vt;L ¼ 12 g/min, C 0ph=460ppm).

0

50

100

150

200

250

300

350

0 50 100 150 200

Time on Stream (min)

T O C

C o n c .

( p p m )

Exit

Inlet

Fig. 4. Extent of homogeneous (non-catalytic) reaction in

trickle bed reactor, measured in terms of TOC reduction.




increase due to the heat of reaction on the pellet scale,

therefore the system was considered isothermal.

Eqs. (7)–(11) were solved simultaneously, using

Eqs. (1), (5) and (6), and by the application of

orthogonal collocation on finite elements (OCFE). Five

hundred finite elements were used with 4 collocation

points used for the gas and liquid phase and 6collocation points used for the solid phase [31,40].

7. Mathematical modeling: TOC equations

The mass balance equations for the TOC system were:

liquid phase:

hLvt;LrbedC 0TOC;L

W total

AplanarC planarTOC

þ3 fDeff ;TOCð1 À eSÞC 0TOC;L

R2

Â AsphericalC sphericalTOC

r¼R

¼ 0; ð12Þ

hLvt;LrbedC 0A;L

W total

AplanarC planarA

þ3 fDeff ;Að1 À eSÞC 0A;L

R2AsphericalC

sphericalA

r¼R

¼ 0; ð13Þ

hLvt;LrbedC 0Oxy;L

W total

AplanarC planarOxy À k La

ÂP 0OxyP 0Ox y

H Oxyrmolar

;

water

À C 0Oxy;LC 0Oxy !þ

3 fDeff ;Oxyð1 À eSÞC 0Oxy;L

R2AsphericalC

sphericalOxy

r¼R

¼ 0;

ð14Þ

gas phase:

hGrbedvt;GP 0Oxy

R1T W total

AplanarP planarOxy

þ k LaP 0OxyP 0Oxy

H Oxyrmolar;water

À C 0Oxy;LC 0Oxy

!¼ 0: ð15Þ

solid phase:

Deff ;TOCC 0TOC

RB sphericalC

sphericalTOC À rS rateTOC ¼ 0; ð16Þ

Deff ;AC 0AR

B sphericalC sphericalA À rS rateA ¼ 0; ð17Þ

Deff ;OxyC 0Oxy

RB sphericalC

sphericalOxy À rS rateOxy ¼ 0; ð18Þ

where rateTOC and rateA are represented analogous to

Eq. (5). For the TOC calculations rateOxy is also

represented by rateTOC since oxygen reaction was

assumed stoichiometrically one to one with TOC.

Similarly Eq. (12)–(18) were solved simultaneously,

using Eqs. (2)–(5) and 6, and by the application of

OCFE.

8. Parameters

The equations solved for modeling trickle bedperformance involved a large number of parameters,

many of which were unknown for this reactor system

and were therefore determined from published correla-

tions.

Molecular diffusivity was calculated using the Siddi-

qi–Lucas [41] method. Liquid and gas holdup were

determined by calculating the liquid fraction of the void

space using a correlation developed by Meng and Chung

[42] and the measured bed porosity. The wetting

efficiency was determined based on a correlation

published by Herskowitz and Smith [43]. Analysis of

the wetting efficiency showed it to be greater then 0.97 atall times, allowing the assumption of complete surface

wetting to be made. The reactor gas–liquid mass transfer

coefficient, k La was determined by averaging the k La’s

as calculated by the correlations proposed by Wild [44]

and Midoux [45]. The Ellman correlation [46] was used

to calculate the pressure drop as a parameter necessary

to the Midox correlation.

9. Comparison of experimental data and model

predictions

Using a Fortran program, Eqs. (7)–(11) were solved

simultaneously along the bed of the reactor with the exit

profiles of each finite element used as the inlet profiles

for each successive element in the reactor. A total of 500

finite elements along the bed of the reactor were used for

each final effluent data point simulation. According to

the model the catalyst pellets frequently became oxygen

deficient towards the center of the pellets, resulting in no

reaction occurring in the oxygen deficient zones.

Hydrolysis reactions, although unlikely, may occur

under such conditions; however, it was assumed in the

model that no reactions occur in the absence of oxygen.

Final calculated effluent conversions are compared tothe experimental values for phenol oxidation in the

trickle bed reactor and can be seen in Figs. 6 and 7.

Similarly overall TOC destruction values for the same

experiments were modeled by Eqs. (12)–(18) and are

shown in Figs. 8 and 9.

As can be seen in Figs. 6–9, the conversion was a

function of temperature as well as space time, where

space time was calculated as g-catalyst/(g-phenol/min).

By using this space time definition it was possible to vary

space time without adjusting the mass flowrate of

the aqueous feed, but by adjusting the concentration

of the feed. On each chart, as space time increases the




conversion of phenol and TOC also increases. As

expected, increasing temperature (by moving from

Figs. 6–9) the reactant conversion also increases. In

Fig. 6, for example, it can be seen that two of the

conditions tested were repeat experiments and two

different experimental data points are noted for each

condition. As expected, the two experimental findings

for reaction conversions are in close agreement, but do

not exactly overlap. It is typically expected however that

the predicted data points from the model would

precisely overlap each other for the same condition,but in this case they do not. The reason for this is due to

the catalyst bed deactivating with time and this

deactivation, while not indicated on the two-dimen-

sional chart, is accounted for in the reactor model. This

result also applies to Figs. 7–9.

10. Summary and conclusions

Based on batch reactor experiments a kinetic model

was developed to model the catalytic reaction for phenol

oxidation, as well as another model for overall TOC

destruction. A trickle bed reactor was operated with

pelletized catalyst in order to generate performance data

from a three-phase flow reactor. The technology was

able to achieve up to 80% TOC destruction in the tested

configuration. Predicted results are then compared with

the experimental data.

The complex three-phase reactor was successfully

modeled without the use of data generated in the flow

reactor (with the exception of catalyst deactivation).

Predictions for phenol oxidation using batch derived

kinetics were successfully applied to a complex three-

phase trickle bed reactor over a temperature range from

150–2051C and a space velocity range from 150 to 650 g-cat/(g-phenol/min). Additionally a lumped TOC model

was developed that successfully predicted batch as well

as flow reactor performance. It can be concluded that

the three-phase trickle bed reactor can be modeled based

on batch derived kinetic data and established phase

behavior correlations.

Acknowledgements

This project has been funded by Grants 104TAM2358

and 026TAM3358 in part with Federal Funds as part of

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

150 170 190 210 230 250 270 290

Space Time (g-cat/(g-phenol/min))

C o n

v e r s i o n

Model

Actual

Fig. 6. Phenol conversion in a trickle bed reactor at 203–2071C.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.70.8

0.9

1

50 100 150 200 250 300 350 400

Space time (g-cat/(g-phenol/min))

C o n v e r s i o n

Model

Actual

Fig. 7. Phenol conversion in a trickle bed reactor at 193–1971C.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

150 200 250 300 350 400 450


C o n v e r s i o n

Model

Actual

Fig. 8. TOC conversion in a trickle bed reactor at 203–2071C.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

0 100 200 300 400 500 600


C o n v

e r s i o n

Model

Actual

Fig. 9. TOC conversion in a trickle bed reactor at 193–1971C.




the program of the Gulf Coast Hazardous Substance

Research Center which is supported under cooperative

agreement R815197 with the United States Environ-

mental Protection Agency and in part with funds from

the State of Texas as part of the program of the Texas

Hazardous Waste Research Center. The contents do not

necessarily reflect the views and policies of the US EPAor the State of Texas nor does the mention of trade

names or commercial product constitute endorsement or

recommendation for use.

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[4] Copa WM, Momont JA. In: H. M. Freeman, editor.

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