FINAL TECHNICAL REPORT March 1, 2013, through July 31 ...

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FINAL TECHNICAL REPORT

March 1, 2013, through July 31, 2014

Project Title: ON-SITE PRODUCTION OF LIME TO REMOVE SO3 AND HCL

FROM ILLINOIS COAL COMBUSTION FLUE GAS

ICCI Project Number: 13/6A-1

Principal Investigator: Hong Lu, ISGS/UIUC

Project Manager: Debalina Dasgupta, ICCI

ABSTRACT

The project aimed at producing highly reactive calcium-based sorbents using a batch

reactor and a 5-10 lbs/hr sorbent activation process (SAP) unit at the Illinois State

Geological Survey (ISGS) for the removal of sulfur trioxide (SO3) and hydrochloric acid

(HCl) in Illinois coal combustion flue gases. More than twenty sorbents were prepared

and the reactivities of these measured using a thermogravimetric analyzer (TGA) and a

tubular reactor. SAP-limes performed better for HCl capture than commercial Trona and

quicklime but exhibited a lower capacity/reactivity for SO3 and sulfur dioxide (SO2)

removal. Partially calcined limestone had the maximum capacity for HCl removal at 70%

extent of calcination (EOC) and sorbents doped with calcium acetate hydrate

[(Ca(OAc)2·H2O)] had a greater than 30% higher capacity for HCl capture.

A techno-economic analysis for a SAP unit installed in a 550 MWe power plant burning

a high-Cl, high-S Illinois No. 6 coal was also conducted. The analysis assumed 70% EOC

during the sorbent production, 40% sorbent utilization during the acid gas removal, and a

90% HCl removal rate. The results revealed that the SAP technology is cost-effective for

on-site production of lime sorbents for HCl and SO3 removal. SAP’s overall cost of acid

gas removal is about 37% lower than a conventional dry sorbent injection (DSI) process.

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EXECUTIVE SUMMARY

The U.S. EPA’s final standards for mercury and other air toxics emissions from electric

generating units (EGUs), commonly referred to as the “Mercury and Air Toxics

Standards” (MATS) set emission limitations on hydrogen chloride (HCl) at 0.02 and

0.0004 lb HCl/MWh for coal-fired existing and new EGUs, respectively. Although HCl

in the flue gas can be efficiently captured in a flue gas desulfurization (WFGD) process

to meet the MATS’ requirement, there is no effective and economic approach to remove

the accumulated chlorides from the scrubber liquid. The application of WFGD facilities

in the utility industry to reduce sulfur dioxide (SO2) emission also leads to increasing

concerns on plume opacity caused by sulfur trioxide (SO3).

Dry sorbent injection (DSI) processes (Figure 1) using sorbents such as hydrated lime

and Trona are often employed to mitigate SO3, SO2, and HCl. The advantages of a DSI

process over a WFGD are: 1) lower capital and raw materials costs; 2) smaller

installation foot print; 3) ease of operation; and 4) more flexibility to fuel changes. A DSI

process upstream of the plant’s particulate control device (PCD), such as a baghouse, can

capture SO3 and HCl in the flue gas upstream of a WFDG. Sorbents may be injected at

different locations between the plant’s economizer and the PCD. Co-benefits of the

removal of HCl and SO3 include: 1) less equipment fouling and corrosion; 2) higher plant

efficiency by lowing air preheater operation temperature; and 3) higher efficiency for

mercury removal by activated carbon injection (ACI).

Figure 1. Block diagram of a SAP or DSI for acid gas removal in a typical power plant.

The Illinois State Geological Survey (ISGS) of the University of Illinois at Urbana-

Champaign, and Electric Power Research Institute (EPRI) have developed a patented

technology, sorbent activation process (SAP), for on-site production of sorbents to

removal pollutants in coal combustion flue gas. SAP can be used as a simple alternate to

DSI for pollutant mitigation. SAP technology allows for producing an acid gas sorbent at

a utility site and directly injecting the fresh sorbent into the flue gas upstream of the

plant’s PCD.

The project aimed at producing highly reactive calcium-based sorbents using a batch

reactor and the 5-10 lbs/hr SAP unit at the ISGS for the removal of SO3 and HCl in Illinoi

coal combustion flue gases. More than twenty sorbents were prepared and the reactivities

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of the sorbents were measured using a thermogravimetric analyzer (TGA) and a tubular

reactor. Performance and properties of the batch sorbents provided useful guidance to

prepare SAP sorbents. Sorbent characterizations included extent of calcination (EOC),

nitrogen adsorption for surface area, and particle size distribution. ISGS prepared lime

sorbents exhibited more than twice of the capacity of HCl capture in comparison to Trona

and quicklime at the same testing conditions, but exhibited a lower performance for SO3

and SO2 removal. The lab-synthesized sorbents had the best performance for HCl

removal when the EOC of the sorbents was at about 70%. Sorbents doped with calcium

acetate hydrate [Ca(OAc)2·H2O] had more than 30% higher capacity for HCl capture. A

combination of various sorbents may maximize HCl/SO3/SO2 removal by taking

advantage of the sorbents’ preferential adsorption on individual gases.

A techno-economic analysis study of SAP technology was performed for producing lime

sorbents for an on-site unit installed in a 550 MWe power plant burning a high-Cl, high-S

Illinois No. 6 coal. The analysis assumed producing a 70% EOC lime sorbent, 40%

sorbent utilization during the acid gas removal, and a 90% HCl removal rate. The results

from this techno-economic analysis revealed that SAP is a cost-effective technology for

producing lime sorbents for HCl and SO3 removal. SAP’s overall cost of acid gas

removal is about 37% lower than a conventional DSI process.

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OBJECTIVES

The project aimed at producing highly reactive calcium-based sorbents using a batch

reactor and the bench SAP unit at the ISGS for the removal of SO3 and HCl in Illinois

coal combustion flue gases. Specific objectives of the project include: 1) the

thermodynamic analysis of the reactions between selected sorbents and HCl/SO2/SO3 at

various process conditions; 2) the preparation of sorbents using a batch reactor and the

bench-scale SAP unit; 3) the evaluation of SAP sorbents using a TGA and a tubular

reactor in a typical Illinois coal flue gas; and 4) the techno-economic analysis of the on-

site sorbent production process installed in a 550 MWe power plant burning a high-S,

high-Cl Illinois coal.

The project tasks were as follows.

Task 1. Sorbent Thermodynamic Analysis A thermodynamic study was performed to identify operating conditions for achieving 90%

HCl removal. Corresponding equilibrium curves (partial pressure vs. temperature) were

developed for reactions between the sorbent and HCl, SO2, and SO3.

Task 2. Sorbent Preparation Using a Bench-Scale SAP

Over 20 sorbents were prepared using a batch reactor and the 5-10 lbs/hr bench-scale

SAP unit at the ISGS. Selected samples were characterized using various techniques such

as nitrogen adsorption for Brunauer-Emmett-Teller (BET) surface area and laser

scattering for particles size distribution.

Task 3. Sorbent Evaluation

The sorbents prepared in Task 2 were evaluated using a TGA and a tubular reactor with

acid gases in a typical Illinois coal flue gas. The temperature window to be examined in

the tests was based on the results obtained from the thermodynamic analysis in Task 1.

Task 4. Techno-Economic Analysis

A techno-economic analysis of the on-site sorbent production process for the removal of

SO3 and HCl in a 550 MWe power plant burning a high-S, high-Cl Illinois coal was

conducted. Capital and operating & maintenance (O&M) costs associated with the on-site

sorbent production and injection system were individually calculated.

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INTRODUCTION AND BACKGROUND

HCl is a hazardous atmospheric pollutant which has adverse effects on both human health

and the wider environment. The U.S. EPA restricts HCl emission to 0.02 and 0.0004 lb

HCl/MWh for existing and new EGUs respectively. [1] In Illinois and other Midwest

bituminous coals, the chlorine content can be as high as 0.54 wt% [2], well above 0.1 wt%

average chlorine content of all U.S. coals. Although wet scrubbers are primarily used for

the removal of SO2, they can effectively capture HCl (>95%) to satisfy the EPA’s

regulations on HCl emissions. However, the removal of HCl via a WFGD leads to

substantial accumulation of chlorides in the scrubber liquid, causing severe equipment

corrosion issues and extra cost to periodically purge the system with low chloride water.

A cost-effective approach to remove HCl upstream of a scrubber is necessary for abiding

by the EPA regulations and minimizing corrosion to the system.

The installation of a NOx removal unit and a WFGD allows a power plant to burn high-S,

high-Cl Illinois coal. The SO3 can only be partially captured at a particulate control

device and a wet scrubber. SO3 in the flue gas may form sub-micro aerosols, causing

public concerns on a visible plume. SO3 mitigation also improves mercury removal by

use of activated carbon injection [3] with a subsequent increase of energy efficiency as a

result of a lower air preheater operation temperature.

DSI is a common technology for the removal of SOx and HCl. (Figure 1) Compared to a

scrubbing system, this technology features 1) lower capital cost; 2) smaller foot print; 3)

user friendly operation; 4) lower raw materials cost; and 5) fuel flexibility. DSI is mostly

used as an alternative to SO2 removal technology for low sulfur coal combustion flue

gases. In a power plant application, the SAP can be used as an alternative to a DSI

system, combining on-site sorbent production from raw materials and sorbent injection

for pollutant mitigation.

In this project, we used a batch reactor and a bench scale SAP unit to produce highly

reactive lime sorbents and evaluated performance of these sorbents for the removal of

acid gases from flue gas at a desired temperature window selected through

thermodynamic analysis. The concentrations of acid gases are typical of a flue gas

generated from burning Illinois basin coal. A techno-economic analysis was conducted

for an on-site sorbent production process for acid gas removal installed in a 550 MWe

power plant burning a high-S, high-Cl Illinois No. 6 coal.

EXPERIMENTAL PROEDURES

Sorbent Preparation

More than twenty sorbents were prepared using a batch reactor (Tables 1 and 2) and the

bench scale SAP unit (Table 3). The performance and properties of the batch reactor-

made sorbents provided a guideline to the sorbent preparation using the SAP, a more

complicated unit than the batch reactor. A typical procedure using the batch reactor is as

follows:

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About one gram of the limestone was evenly placed on a 1.5”x2” flat metal sample boat.

The sample boat was inserted in the cold section of a horizontal quartz tube (2-in ID)

which was preheated to the reaction temperature by a Lindberg furnace. A metal boat was

chosen to facilitate heat transfer during the calcination. The quartz tube was purged with

N2 for 10 min before the sample boat was pushed to the hot zone of the furnace and kept

at this location for a predetermined time. The temperature was controlled and monitored

by a K-type thermocouple above the sample boat. A N2 flow of 250 sccm purged the

reactor during the entire process. At the end of the reaction time, the sample boat was

brought back to the cool section of the reactor for a few minutes and stored in a container

with desiccant. During this procedure, limestone was partially decomposed to CaO

(CaCO3 CaO + CO2). For sorbents with dopants, the dopants (calcium acetate hydrate

sodium acetate , sodium chloride, sodium carbonate) were added onto CaO/CaCO3

sorbent via the following approach: 1) Two grams of limestone and calculated amount of

dopants were added into 50 ml DI water and the slurry was vigorously stirred for 10

minutes; 2) The slurry was evaporated to a paste in the air and then dried in an oven at

105 °C overnight; 3) The solids were calcined at 900 °C in the presence of nitrogen.

Products obtained were sealed in bottles filled with nitrogen. The batch sorbents have

been summarized in Tables 1 and 2. Note that the BET surface areas in the tables are

based on the total sorbent weight, not the CaO content of the sorbents.

Table 1. Preparation conditions and properties of the batch sorbents.

Sample Temperature Residence

time EOC

BET surface

area Note

°C min % m2/g

IF001 1000 1.5 31 3.6 LS-AC3

IF002 1000 3 92 7.6 LS-AC3

IF003 900 3 45 4.9 LS-AC3

IF004 900 5 91 8.6 LS-AC3

IF005 800 3 11 2.4 LS-AC3

IF006 800 5 20 3.2 LS-AC3

IS020 900 5 86 10.7 LS-X8

IS021 900 10 98 9.2 LS-X8

IS022 900 15 97 10.7 LS-X8

Table 2. Preparation conditions and properties of the batch sorbents with dopants.

Sample Temperature Residence

time EOC Note

°C min %

IS008 900 3 43 2wt% NaCl/LS-AC3

IS013 900 3 23 2wt% Ca(OAc)2 /LS-AC3

IS014 900 3 66 2wt% NaOAc/LS-AC3

IS018 900 3 42 2wt% Na2CO3/LS-AC3

IS019 900 3 49 2wt% Na2SO3/LS-AC3

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The bench-scale SAP (Figure 2) is an L-shaped entrained-flow reactor with an inner

diameter of 6 inch. It has several ports for either monitoring the temperatures or injecting

limestone particles. A Krom Schroder BIC-65 burner, a pre-mix burner with a maximum

capacity of 70 kW is located about 22-inch upstream of the injection port and burnt

propane as the main heat source in this unit. After preheating the SAP, limestone was

injected into the reactor via a volumetric screw feeder that aided by a purge gas. Inside

the SAP, limestone was partially decomposed to CaO. Some of the products are expected

to be re-carbonated (CaO + CO2 CaCO3) in the tail gas before it can be collected at the

effluent of the SAP through a cyclone. Doped sorbents were produced by feeding

premixed limestone and dopant. Detailed preparation conditions and properties of the

SAP-made sorbents are given in Tables 3 and 4.

Figure 2. Front look of SAP and temperature port locations.

Table 3. Preparation conditions and properties of the SAP sorbents.

Sample

SAP operating temperatures Sorbent properties

Note T1 T2 T3 EOC

BET surface

area

°C °C °C % m2/g

IS010 677 581 464 13 2.0 LS-AC3

R11 980 768 629 61 5.0 LS-X8

R17 1044 813 648 67 7.3 LS-X8

R13 1110 886 727 81 5.9 LS-X8

R15 1146 944 773 86 4.1 LS-X8

R21 1137 914 752 86 4.0 2 wt.% Ca(OAc)2

/LS-X8

R31 1174 970 793 88 3.2 2 wt.% Na2SO3

/LS-X8

M11 936 681 526 67 4.2 LS-M6

M12 829 610 493 51 3.3 LS-M6

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Table 4. Median particle size of raw limestone and SAP sorbents.

Sample Median particle size

µm

LS-X8 18

R11 20

R17 21

R13 23

R15 24

R21 21

R31 24

LS-M6 9

M12 12

Sorbent Characterizations

Selected samples were subjected to various characterizations. BET surface area was

measured by N2 adsorption and desorption isotherms at -196 °C using a Monosorb

surface area analyzer (Quantachrome Inc.). Volume-based particle size distribution

(PSD) was measured using a laser scattering particle size analyzer (Horiba Instruments

Inc., LA-950). The instrument can measure particles ranging from 0.3 nm to 30 mm. The

sorbents were also subjected to full calcination using a TGA (Thermo Scientific

Versatherm) to determine the EOC of a sorbent.

Sorbent Evaluation

The performance of the sorbents for acid gas adsorption was evaluated using the TGA

and a tubular reactor. TAG is a common and effective tool to evaluate the performance of

lime sorbents for acid gases (HCl and/or SO2) [4-6] and CO2 capture [7-9] despite that

the reaction conditions including reaction time and gas-solid contact patterns are much

different in the TGA from a real applications. A typical procedure using the TGA for the

sorbent evaluation was as follows: 1) a ~4 mg of sample was heated from room

temperature to reaction temperature in N2 at 10 °C/min; 2) the sample was kept at

reaction temperature for 10 min; 3) N2 was switched to reacting gas such as 600 ppm

HCl, 3000 ppm SO2, or 105 ppm SO3 (balance N2) at 400 °C for 90 min unless otherwise

stated.

Performance of a selected sorbent and a commercial sorbent were evaluated using the

tubular reactor (1/4-in ID) with typical concentrations of acid gases in an Illinois coal

burning flue gas: 300 ppm HCl, 2500 ppm SO2, and 50 ppm SO3. The system included a

gas mixing section and a tubular reactor section. The gas mixing section included gas

cylinders and mass flow controllers (MFC) to generate a desired gas stream. SO3 was fed

into the gas line using a Perkin Elmer syringe pump. The gas line was heated with a

heating tape to prevent the SO3 from condensing. The reactor section included a tubular

reactor, a furnace, and a temperature controller (Omega CN7500). Gas breakthrough

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curves were obtained via monitoring gas concentrations in the effluent using a mass

spectrometry technique-based residue gas analyzer (MKS Cirrus 2). To enhance the

utilization of the sorbent, inert silica sand (210 to 297 µm) was used to disperse sorbent

particles at a ratio of 20 g-silica/g-sorbent. Roughly, the residence time of the reaction

gas within the sorbent bed was extended for 20 times. Moreover, channeling effect was

also effectively reduced, improving better gas-particle contact effect.

RESULTS AND DISCUSSION

Thermodynamic Analysis

Thermodynamic analysis methodology

The thermodynamic analysis study was focused on reactions between CaO and acid

gases. Dominant products of a reaction system at equilibrium were identified by

minimizing the change of Gibbs free energy of the reaction system according to Eq. 1:

ΔG = ∑ (µio + RTlnϕiPi) + ∑ (µj

o + RTlnaj) (Eq. 1)

where ΔG is the change of Gibbs free energy, i is a gas component, j is a pure condensed-

phase component, n is the number of moles, µo is the standard chemical potential, ϕ is the

fugacity coefficient, P is pressure, T is temperature, R is the gas constant, and a is the

activity of the reaction species. The equilibrium constant K of a reaction was calculated

according to Eq. 2:

K = exp[–ΔG0/(RT)] (Eq. 2)

Equilibrium pressures of a reaction at various temperatures were calculated using Eq. 3:

Kp = ∏ Pia (Eq. 3)

where Kp is the equilibrium constant of a reaction and a is the stoichiometric number of

gas component i.

Adsorption reaction

The dominate reactions between a CaO sorbent and acid gases identified from the

thermodynamic analysis are as follows (100-900 °C):

CaOHCl (s) = CaO (s) + HCl (g) (Rx. 1)

CaCl2 (s) + H2O (g) = CaO (s) + 2HCl (g) (Rx. 2)

CaSO3 (s) = CaO (s) + SO2 (g) (Rx. 3)

CaSO4 (s) = CaO (s) + SO3 (g) (Rx. 4)

The equilibrium pressure as a function of temperature obtained via Eq. 3 is given in

Figure 3. Above each equilibrium curve, the adsorption reactions occur and the acid gas

is captured by the sorbent. Below each equilibrium curve, the desorption reactions

(reverse reaction) proceed.

Temperature window

In an ideal operation window, a sorbent reacts with both HCl and SO3 but not SO2 since

SO2 can be captured by a usually less expensive approach of flue gas desulfurization

process. When the temperature is lower than 790 °C, the tendency to react with the acid

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gases by CaO is HCl<SO2<SO3 thermodynamically as shown in Figure 3. At

temperatures above 790 °C, forward reactions of Rx. 1 and 3 and reverse reactions of Rx.

2 and 4 occur, leading to HCl and SO3 but not SO2 capture by the sorbents.

Figure 3. Equilibrium pressures of acid gases using a CaO sorbent.

Figure 4. Adsorption temperature window for a 90% HCl removal from a 600 ppm HCl

stream using a CaO sorbent. Left: Rx. 1 mechanism; Right: Rx. 2 mechanism.

To capture 90% of HCl from a flue gas containing 600 ppm HCl, the thermodynamic

analysis suggests that the adsorption operation should occur below 510 °C (Figure 4).

Selectivity

The selectivity of HCl or SO3 over SO2 by a lime sorbent is defined as the ratio of

equilibrium pressure of HCl (PHCl/PSO2) or SO3 (P SO3/PSO2) over SO2. The lower the ratio

is, the better the sorbent adsorbs HCl or SO3. Figure 5 exhibits that regardless of which

HCl adsorption mechanism dominates, a higher reaction temperature leads to a lower

ratio of the equilibrium pressures of HCl over SO2, suggesting that the preferential

capture of HCl over SO2. With decreasing reaction temperature, SO3 adsorption is

preferred over SO2 (Figure 6). In this project, major tests of the sorbents were run at

400 °C to keep high selectivity of HCl over SO2.

200 400 600 8001E-14

1E-11

1E-8

1E-5

0.01

790 C

Ac

id g

as

eq

uilib

riu

m p

res

su

re (

ba

r)

T (C)

CaOHCl=CaO+HCl

CaCl2+H2O=CaO+2HCl

CaSO3=CaO+SO2

CaSO4=CaO+SO3

200 400 600 8001E-14

1E-11

1E-8

1E-5

0.01600 ppm

Desorption

HC

l e

qu

ilib

riu

m p

res

su

re (

ba

r)

T (C)

CaOHCl=CaO+HCl

Adsorption

60 ppm

520 C

200 400 600 8001E-14

1E-11

1E-8

1E-5

0.01600 ppm

Desorption

HC

l e

qu

ilib

riu

m p

res

su

re (

ba

r)

T (C)

CaCl2+H2O=CaO+2HCl

Adsorption

60 ppm

510 C

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Figure 5. Adsorption selectivity of HCl over SO2 using a CaO sorbent. Left: Rx. 1

mechanism; Right: Rx. 2 mechanism.

Figure 6. Selective adsorption of SO3 over SO2 using a CaO sorbent.

Sorbent Preparation and Characterizations

Particle size

Particle size measurement shows that the median particle size of limestone AC-3 (IF000)

was 4.0 µm and 90% (volume) of the particles were smaller than 6.7 µm (Figure 7).

Small sizes are beneficial for fast calcination because shorter time exposure to high

temperature is required during calcination. High temperature calcination is known to lead

to particle sintering and performance deterioration. The results in Figure 7 show that the

produced sorbents had larger size than the raw limestone (IF000). There are two possible

explanations: particle sintering due to high temperature and expanding due to the fast

release of CO2 during the calcination. The median size of sample IF002 (9.1 µm) was

about 60% larger than that of sample IF004 (5.7 µm) even though these samples had

comparable EOCs. The larger size of the former sample was probably caused by sintering

at the higher preparation temperature.

200 400 600 8001

100

10000

1000000

1E8

PH

Cl/P

SO

2

T (C)

2CaO+HCl+SO2

200 400 600 8000.1

10

1000

100000

1E7

1E9

1E11

PH

Cl/P

SO

2

T (C)

2CaO+2HCl+SO2

200 400 600 800

1E-17

1E-14

1E-11

1E-8

PS

O3

/PS

O2

T (C)

2CaO+SO3+SO2

12

Figure 7. Particle size distributions of the limestone (IF000) and sorbents. Left: Size

distribution vs. diameter; Right: Accumulated size distribution vs. diameter.

The sizes of the SAP sorbents exhibited the similar trend: particles were bigger when the

sorbents were prepared at higher temperature (Figure 8). Median particle sizes of the

sorbents from LS-X8 without dopants increased linearly with the preparation temperature

(Figure 8). The median particle size of the sorbent R15 (prepared at 944 °C, T2) was 24

µm, which is 25% greater than that of the sorbent R11 (prepared at 768 °C, T2). The

median size of sorbent R15 was also 33% greater than that of the raw limestone at 18 µm.

Particles tend to sinter when temperature is higher than their Tammann temperature,

which is estimated as 0.52 times of the particles’ melting point temperature (K). The

higher temperature caused sintering which led to larger particle sizes.

Figure 8. Median particle size and EOC of the SAP-sorbents (R11, R17, R13, and R15)

made at different temperatures with the same limestone.

Dopants also affected particle sizes. Sorbents R15, R21, and R31 were prepared at similar

SAP conditions but with different dopants. The sorbents had similar EOCs (86% - 88%).

However, they showed different size distributions and their median sizes ranged from 21

to 24 µm (Table 4 and Figure 9). Median size of sorbent R21 was one eighth smaller than

1 100

5

10

15

20

q (

%)

Diameter (m)

IF000

IF001

IF002

IF003

IF004

IF005

IF006

1 100

20

40

60

80

100

Un

ders

ize (

%)

Diameter (m)

IF000

IF001

IF002

IF003

IF004

IF005

IF006

50%

60%

70%

80%

90%

100%

17

19

21

23

25

700 800 900 1000

EOC

Med

ian

Siz

e (µ

m)

Temperature (°C)

13

those of sorbents R15 and R31. The smaller particles of sorbent R21 might be caused by

the interactive effect between dopant Ca(OAc)2 and the lime particles when the dopant

decomposed during its multiple-step decomposition to CaO.

EOC

EOC is an index of the degree that a raw material is calcined. During the sorbent

preparation, the limestone was partially decomposed to CaO. The sorbents were a

mixture of un-calcined limestone (CaCO3) and freshly produced CaO. To determine the

EOC, the sorbents were subjected to a full calcination using the TGA. The EOC was

calculated from the weight loss during the fully calcination using the following equation:

EOC = (1-wL/wL0)/(1-wL) (Eq. 4)

where wL and wL0 are weight losses during the fully calcination of the sorbent and the

raw limestone, respectively, measured in the TGA calcination. Higher temperature and

longer residence time during the sorbent preparation resulted in higher EOC (Tables 1

and 3). At higher temperature, the limestone had faster kinetics and stronger tendency to

decompose, resulting in the greater EOC. For examples, the batch samples IF002 and

IF004 had EOCs of 92% and 91%, respectively, much greater than those of the samples

IF005 and IF006 (Table 1). The only difference among the four samples is that the former

two were prepared at higher temperatures. Figure 8 demonstrates that the EOC of the

SAP sorbents increased proportionally with the sorbent preparation temperature.

Figure 9. Particle size distributions of SAP-sorbents with different dopants prepared

from LS-X8.

BET surface area

Limestone partially decomposed to CaO during the preparation, releasing CO2 and

producing porous structure that increased the surface area of the sorbents. Surface area of

six batch samples (IF001-IF006) was plot against their EOC. The surface area of the

sorbents increased linearly with the increasing EOC (Figure 10).

1 10 100 1000

0

4

8

12

16

q (

%)

Diameter (m)

R15

R21

R31

14

Figure 10. BET surface area vs. EOC (Batch sorbents from LS-AC3).

Sorbent Evaluation

Sorbent reactivity for HCl adsorption using the TGA

Thermodynamic analysis discussed previously implies that the adsorption should be

carried out at temperatures lower than 510 °C to achieve 90% HCl removal by CaO in a

600 ppm HCl stream and higher temperature is beneficial to selectivity of HCl over SO2.

HCl adsorption capacities of the batch sorbents without dopants are summarized in

Figure 11. IF003 exhibited the largest capacity while IF005 and IF006 showed much

smaller capacities than the other sorbents. The smaller capacities of the latter samples

might be due to their limited surface area accessible by HCl molecules. Although IF002

and IF004 possessed much greater surface area than samples IF001 and IF003, the latter

demonstrated better performance. During sorbent preparation, IF001 and IF003 were

exposed to high temperatures at a shorter period than IF002 and IF004 at 1000 and

900 °C. A greater surface area of a sorbent was due to the formation of porous structures

during CaCO3 calcination. The smaller surface area of samples IF001 and IF003 was due

to the smaller degree of calcination. Limestone decomposition happens gradually from

exterior to interior according to shrinking core model. Partial calcination of IF001 and

IF003 should occur majorly at the exterior of the particles. Therefore, their surface area

might be more concentrated on the exterior part of the particles. IF002 and IF004

developed with greater surface area due to their longer calcination time which calcined

over 90% of the precursor particles. However, during the diffusion controlled reaction,

HCl molecules were consumed up at the exterior before they were able to reach the

surface CaO deep inside the particles during the 90 min reaction period. This indicates

that only partial of the available surface CaO was utilized for HCl adsorption. Since

residence time of sorbent particles in a duct DSI process is only a few seconds, it is

reasonable to predict that only a limited portion of the surface area will have the

opportunity to react with an acid gas. Therefore, it may not be necessary to fully calcine

limestone to produce a high surface area product. Suitable sorbents that are partially

calcined may achieve comparable performance to a fully calcined one but at lower cost.

0 20 40 60 80 100

2

4

6

8

BE

T s

urf

ace a

rea (

m2/g

)

EOC (%)

15

Figure 12 exhibits a more clear view on how the adsorption behavior of IF004 and IF006

evolved with respect to different surface areas. In the initial 10 min, IF004 (8.6 m2/g) and

IF006 (3.2 m2/g) had comparable adsorption rates, indicating that HCl was reacting with

CaO on the most exterior part of the particles. After that, the reaction rate of IF006

gradually decreased and became negligible after about 120 min, indicating HCl diffused

through the CaCl2/CaOHCl product layer. IF004 maintained a faster rate during the entire

900 min adsorption period due to its greater surface area. The somehow slower reaction

rate at later stage could be attributed to larger diffusion resistance of HCl into the interior

pores of the sorbent particles.

Sorbents from different types of limestone showed different performance for HCl

adsorption. Figure 13 showed the performance of two batch sorbents from two types of

limestone at the same conditions. IS020 from LS-X8 exhibited high HCl capture capacity

of about 0.20 g-HCl/g-sorbent while IF004 from LS-AC3 only 0.11 g-HCl/g-sorbent.

Figure 14 presented the performance of two sorbents from two types of limestone made

0 30 60 900.00

0.03

0.06

0.09

0.12

1 2

0.000

0.001

Ca

pa

cit

y (

g-H

Cl/

g-s

orb

en

t)

Time (min)

IF 001

IF 002

IF 003

IF 004

IF 005

IF 006

Cap

acit

y (

g-H

Cl/

g-s

orb

en

t)

Time (min)

IF 001

IF 002

IF 003

IF 004

IF 005

IF 006

0 200 400 600 8000.0

0.1

0.2

0.3

0.4

0.5

0 10 20 30 40 50 60 70 80

0.00

0.01

0.02

0.03

0.04

Cap

acit

y (

g-H

Cl/

g-s

orb

en

t)

Time (min)

IF 004

IF 006

Ca

pa

cit

y (

g-H

Cl/

g-s

orb

en

t)

Time (min)

IF 004

IF 006

0 20 40 60 80

0.0

0.1

0.2

Ca

pa

cit

y (

g-H

Cl/g

-so

rbe

nt)

Time (min)

IS-020

IF-004

0 20 40 60 80

0.00

0.02

0.04

0.06

0.08

0.10

Ca

pa

cit

y (

g-H

Cl/g

-so

rbe

nt)

Time (min)

R11

M11

Figure 12. HCl adsorption over

extended time at 400 °C. Figure 11. HCl adsorption of

various sorbents.

Figure 13. HCl adsorption capacity of

sorbents made at the same conditions

from different types of limestone: LS-

X8 (IS020) and LS-AC3 (IF004).

Figure 14. HCl adsorption capacity of

sorbents made at the similar conditions

from different types of limestone: LS-

X8 (R11) and LS-M6 (M11).

16

at similar conditions from the SAP. M11 prepared from LS-M6 had a capacity of 0.088

g-HCl/g-sorbent, much greater than that of R11 despite that the latter had a greater

surface area and similar EOC. The smaller particle sizes of LS-M6-based sorbents may

have contributed to the better performance.

Sorbents with different dopants shows different performance. HCl adsorption reactivity

of the batch sorbents with dopants is summarized in Figure 15. IS014 with dopant

NaOAc exhibited the highest capacity, about 0.097 g-HCl/g-sorbent in the testing,

followed closely by IS013 with dopant Ca(OAc)2 at 0.075 g-HCl/g-sorbent. Sorbents

doped with Na2SO3 (IS019), NaCl (IS015), and Na2CO3 (IS018) had capacities between

0.03 and 0.55 g-HCl/g-sorbent. It is interesting to note that both IS014 and IS013 were

doped with acetate salts. Compared to the sorbents IS014 - IS019, IS013 had the lowest

EOC (23%). The multiple decomposition steps of calcium acetate hydrate

Ca(CH3COO)2·H2O might have produced porous structure that was beneficial to the gas-

solid reaction [10].

Figure 15. HCl adsorption capacity of sorbents with dopants made from the batch

(Left) and SAP (Right)

In Figure 15, the performance of SAP sorbent (R15) without dopant is compared with the

sorbents (R21 and R31) with dopants that show good performance from batch sorbents.

Doped with Ca(OAc)2 and Na2SO3, R21 and R31 exhibited 31% and 36%, respectively,

higher capacities than R15 for HCl adsorption. Note that all three sorbents were prepared

at similar SAP conditions but R15 did not have any additives. The results demonstrate

that dopants Ca(OAc)2 and Na2SO3 effectively enhanced performance of lime sorbents

for HCl capture. For comparison, the performance of commercial Trona (EnProveTR,

Natronx) for HCl adsorption was also included. The dopant enhanced SAP sorbents

exhibited much better performance than the commercial Trona at the test conditions

either at the initial reaction period or the entire extended reaction stage, indicating the

advantage of SAP sorbents over the Trona for HCl mitigation.

An interesting observation is that a maximum exists when sorbent capacity for HCl

removal vs. EOC is plotted (Figure 16). At higher preparation temperature, the sorbent

had higher EOC as shown in Figure 8, which means that the sorbent possessed greater

percentage of CaO. However, the capacity of the sorbent for capturing HCl did not

always increase with the EOC. There was a maximum capacity for HCl capture at about

70% EOC for both the batch and SAP sorbents. When the EOC was greater than about

0 20 40 60 80

0.00

0.02

0.04

0.06

0.08

0.10

Ca

pa

cit

y (

g-H

Cl/g

-so

rbe

nt)

Time (min)

IS-008

IS-013

IS-014

IS-018

IS-019

0 20 40 60 80

0.00

0.02

0.04

0.06

0.08

0.10

0.12

Ca

pa

cit

y (

g-H

Cl/

g-s

orb

en

t)

Time (min)

R15

R21

R31

Trona

17

70%, the capacity started to decrease with the EOC. Sorbent prepared at higher

temperature or longer residence would sinter more as discussed above. Despite a sorbent

may have a greater EOC (more CaO per gram sorbent), the amount of active CaO sites

accessible by the gaseous HCl may be less than other sorbents at smaller EOCs. The

existing of a maximum capacity is useful to determine optimal conditions for producing

sorbent using the SAP. Energy can be saved by partially calcining limestone while

achieving better performance for acid gas removal. Assuming limestone are sphere

particles, at 70% EOC, only the exterior 1/3 (by diameter) of the limestone particles was

calcined. (Figure 16)

Figure 16. HCl adsorption capacity vs. EOC (Left, batch sorbents IF001-006; Right ,

SAP sorbents R11, 17, 13, and 15).

The HCl capacities of a SAP sorbent (R17) and the commercial Trona were compared at

different testing temperatures (Figure 17). R17 had capacities of 0.014, 0.072, and 0.127

g-HCl/g-sorbent at 200, 300, and 400 °C. The Trona showed very limited capacity for

HCl adsorption at all of the testing temperatures (<0.02 g-HCl/g-sorbent), demonstrating

that the SAP lime sorbents are better than the Trona for HCl adsorption, especially when

the temperature is at 300 °C or higher.

Figure 17. HCl adsorption capacity of a SAP sorbent (R17, Left) and the commercial

Trona (Right) at 200, 300, and 400 °C.

0

0.04

0.08

0.12

10% 30% 50% 70% 90% 110%

Cap

acit

y (g

-HC

l/g-

sorb

en

t)

EOC

0.00

0.04

0.08

0.12

50% 60% 70% 80% 90%

Ca

pa

cit

y (

g-H

Cl/

g-s

orb

en

t)

EOC

0 30 60 900.00

0.03

0.06

0.09

0.12

Ca

pa

cit

y (

g-H

Cl/

g-s

orb

en

t)

Time (min)

400 C

300 C

200 C

0 20 40 60 80

0.00

0.01

0.02

0.03

0.04

0.05

Ca

pa

cit

y (

g-H

Cl/g

-so

rbe

nt)

Time (min)

400 oC

300 oC

200 oC

Limestone particles calcined

(orange) at 70% EOC

18

Sorbent reactivity for SO2 adsorption using the TGA

Selected sorbents were tested for SO2 capture. The sorbents IS020 - IS022 showed

capacities between 0.065 and 0.070 g-SO2/g-sorbent (Figure 18). All of the three sorbents

displayed two distinguished reaction states: a fast reaction in the initial 15 minutes

followed by a much slower one. The Trona, however, exhibited a capacity of 0.106 g-

SO2/g-sorbent at the same conditions and it did not show a slow reaction stage at the

testing conditions (Figure 20). Even at 200 and 300 °C, the Trona captured 0.05 and 0.06

g-SO2/g-sorbent.

The SAP sorbents M11 and M12 also exhibited low capacity for SO2 capture (Figure 20)

while had high capacity for HCl adsorption (Figure 14). Note that the Trona had an SO2

capacity of 0.106 g/g-sorbent at 400 °C (Figure 19), an order greater than those of the

SAP sorbents, and an HCl capacity of 0.016 g/g-sorbent (Figure 17), an order smaller

than those of the SAP sorbents. These suggest that the SAP lime sorbents be effective for

HCl adsorption but not SO2.

0 20 40 60 80

0.00

0.02

0.04

0.06

0.08

0.10

Cap

acit

y (

g-S

O2/g

-so

rben

t)

Time (min)

IS-020

IS-021

IS-022

0 20 40 60 80

0.00

0.03

0.06

0.09

0.12

Ca

pa

cit

y (

g-S

O2/g

-so

rbe

nt)

Time (min)

Trona 300C-200C

Trona 300C-300C

Trona 300C-400C

0 20 40 60 80

0.000

0.004

0.008

0.012

0.016

Ca

pa

cit

y (

g-S

O2/g

-so

rbe

nt)

Time (min)

M12

M11

0 30 60 900.00

0.03

0.06

0.09

0.12

Ca

pa

cit

y (

g/g

-so

rbe

nt)

Time (min)

R17 HCl

R17 SO2

Quicklime HCl

Quicklime SO2

Figure 19. SO2 adsorption capacity

of the Trona at 200, 300, and 400 °C

Figure 18. SO2 adsorption capacity

of sorbents made from LS-X8.

Figure 20. SO2 adsorption capacity of

SAP sorbents from LS-M6.

Figure 21. Comparison of SAP lime

sorbent (R17) with a commercial

quicklime for SO2 and HCl adsorption

19

In Figure 21, a SAP sorbent (R17) without any dopant is compared with commercial

quicklime. While the SAP sorbent exhibited faster reaction for HCl adsorption than the

quicklime, it was slower for SO2 removal. This reaction behavior is wanted since the lime

sorbent was expected to capture HCl, but not SO2, which can be removed by a usually

less expensive approach, a wet scrubber. When both HCl and SO2 are required to be

removed from a gas stream, a combined utilization of the SAP sorbents and a commercial

Trona or quicklime is expected to have a better performance.

Dopants also showed effect on the SO2 adsorption. SO2 adsorption capacities of the batch

sorbents with various dopants are summarized in Figure 22. Sorbent with dopant NaOAc

(IS014) captured 0.038 g-SO2/g-sorbent. Sorbents with dopants of Ca(OAc)2 (IS013),

Na2CO3 (IS018), or Na2SO3 (IS019) captured less than 0.02 g-SO2/g-sorbent. Sorbent

IS013 with dopant Ca(OAc)2 captured 0.075 g-HCl/g-sorbent (Figure 15) while only

moderate SO2, making it a promising candidate to selectively remove HCl over SO2.

Sorbents with dopants of Ca(OAc)2 and Na2SO3 were prepared using the SAP. The

performance of the sorbents was compared with the commercial Trona and the sorbent

R15 which was prepared at similar conditions using the SAP but no dopant (Figure 23).

R31 doped with Na2SO3 had a capacity of 0.016 g-SO2/g-sorbent. R21 with Ca(OAc)2

and R15 without any dopants had capacities of 0.010 and 0.008 g-SO2/g-sorbent,

respectively. Although R21 and R31 had 31% and 36% higher capacities than R15 for

HCl adsorption, they had 25% and 100%, respectively, higher capacities than R15 for

SO2 adsorption. This suggests that dopant Ca(OAc)2 not only enhanced the sorbent’s

capability to capture HCl but also enhanced its selectivity on HCl over SO2. The

enhancement of HCl adsorption over SO2 by R21 (31% vs. 25%) might due to the

structure change caused by Ca(OAc)2 that promotes diffusion of reaction gas to the active

CaO sites. The enhancement SO2 adsorption over HCl by R31 (100% vs. 36%) with

Na2SO3 might due to seeding effect of SO2-

that is preferential to SO2 adsorption,

resulting in a significant enhancement for SO2 adsorption.

0 20 40 60 80

0.00

0.01

0.02

0.03

0.04

0.05

Ca

pa

cit

y (

g-S

O2/g

-so

rbe

nt)

Time (min)

IS-013

IS-014

IS-018

IS-019

0 20 40 60 80

0.000

0.004

0.008

0.012

0.016

0 20 40 60 80

0.000

0.004

0.008

0.012

0.016

Ca

pa

cit

y (

g-S

O2/g

-so

rbe

nt)

Time (min)

R15

R21

R31

0 20 40 60 800.00

0.03

0.06

0.09

0.12

Ca

pacit

y (g

-SO

2/g

-so

rben

t)

Time (min)

Trona

Ca

pa

cit

y (

g-S

O2/g

-so

rbe

nt)

Time (min)

Figure 22. SO2 adsorption capacity of

sorbents with dopants.

Figure 23. SO2 adsorption capacities of

sorbents with dopants: R21 with 2 wt.%

Ca(OAc)2 and R31 with 2 wt.% Na2SO3.

20

Sorbent reactivity for SO3 adsorption using the TGA

A SAP sorbent (IS010) and the Trona were compared for SO3 adsorption. After one hour

of reaction, IS010 showed a capacity of 0.0017 g-SO3/g-sorbent at 400 °C while the

Trona sorbent exhibited a capacity of 0.005 g-SO3/g-sorbent at the same conditions

(Figure 24). Noting that the SAP sorbents showed much better HCl capacity, a combined

utilization of SAP sorbents and a Trona is expected to have a better performance when

both HCl and SO3 are required to be removed from a gas stream.

Figure 24. SO3 adsorption capacity of SAP and Trona sorbents at 400 °C.

Sorbent reactivity for acid gas adsorption using the tubular reactor

A SAP sorbent (R17) and a commercial quicklime sorbent were evaluated with acid

gases typical in an Illinois basin coal combustion flue gas (300 ppm HCl, 2500 ppm SO2,

and 50 ppm SO3 in N2) using the tubular reactor. Though no HCl and SO3 breakthrough

was observed during the 100 min tests, SO2 in the effluent almost reached a plateau

(Figure 25), indicating the sorbents favored reactions with HCl and SO3. An interesting

observation is that the amount of SO2 captured by R17 was about half of that by the

quicklime in both the TGA and tubular reactor tests (Table 5).

Figure 25. Evolution of HCl and SO2 in the effluent of the tubular reactor using acid

gases typical in an Illinois coal combustion flue gas.

0 20 40 60 80

0.00

0.01

0.02

0.03

0.04

0.05

Ca

pa

cit

y (

g/g

-so

rbe

nt)

Time (min)

IS-010 HCl

IS-010 SO3

Trona HCl

Trona SO3

21

Table 5. Sorbent utilization and weight change for HCl and SO2 adsorption tested with

the TGA and tubular reactor.

SO2 HCl

Sorbent utilization using

the tubular reactor

Capacity using

the TGA, g/g

Capacity using the

TGA, g/g (at 70 min)

R17 17% 0.028 0.114

Quicklime 30% 0.063 0.067

Techno-Economic Analysis of SAP for Acid Gas Removal

A techno-economic analysis was conducted for a conceptual on-site SAP process to

remove acid gases in the flue gas of a 550 MWe power plant burning high-Cl, high-S

Illinois No. 6 coal. The acid gas removal cost was compared with that of a traditional DSI

process. Sensitivities of the cost of the SAP to sorbent utilization rate and limestone price

were also addressed.

Process description and assumptions

A schematic diagram of a power plant using either a SAP or DSI system for acid gas

mitigation is presented in Figure 1. Sorbent powders can be injected using either the SAP

or DSI at different locations between the furnace boiler and the inlet of the plant’s

particulate control device. In the SAP technology, lime sorbents are produced on-site and

injected directly into the plant’s flue gas. The process takes advantage of on-site coal and

limestone supplies, handling facilities, and energy sources. Pulverized coal and limestone

are separately injected into the combustion and sorbent activation zones of the SAP. In

the SAP, micronized (<45 microns) limestone particles are thermally decomposed to lime

in 1 to 3 seconds. The freshly produced lime sorbent particles are injected directly into

the flue gas. Essentially, a SAP unit is a combination of a limestone calciner unit and a

DSI unit (Figure 26). Thus, the cost of the SAP system was estimated to be a combined

cost of the two units.

A 550 MW net output subcritical unit (Case 9 in Reference [11]) was selected for the

analysis. Selected properties of an Illinois No. 6 coal used in this case are listed in Table

6. In Case 9, Illinois No. 6 coal and primary air are introduced into the boiler through the

wall-fired burners and the flue gas exits the boiler through the selective catalytic reactor

and is cooled to 169 °C in the air preheater before passing through a baghouse for

particulate removal. Mass balance, flow rates, and assumptions of the sorbent

performance in the SAP are given in Tables 7 and 8. Because no SO3 concentration is

reported in case 9 in [11], it was assumed that 98.5% and 1.5% of the total sulfur in the

coal was converted to SO2 and SO3, respectively, generating a SO2 concentration of 2,102

ppm. The amounts of the limestone and coal feed rates were calculated using the process

conditions and assumptions listed in Tables 6 to 9.

22

Figure 26. SAP (left) and traditional DSI (right) for acid pollutant control.

Table 6. Properties of Illinois No. 6 coal. [11]

Item As Received Dry

HHV, kJ/kg 27,113 30,506

Chlorine, wt.% 0.29 0.33

Sulfur, wt.% 2.51 2.82

Ash 9.70% 10.91%

Table 7. Mass balance and flow rate. [11]

Item Value Unit

Coal feed rate 198,391 kg-coal/hr

Volumetric flue gas rate before the baghouse 72,904 kg-mol/hr

Cl flow rate before the baghouse 575 kg-Cl/hr

S flow rate before the baghouse 4,980 kg-S/hr

HCl concentration before the baghouse 222 ppmv

SO2 concentration before the baghouse 2,102 ppmv

SO3 concentration before the baghouse 32 ppmv

Fly ash rate 15,390 kg-ash/hr

Table 8. Mass balance and assumptions for SAP.

Item Value Unit

Sorbent utilization by molar 40%

HCl removal rate 90%

SO3 and SO2 removal rate 40%

Limestone feed rate 17,385 kg-LS/hr

Limestone calcination rate by the SAP 70%

Total sorbent produced by the SAP 12,030 kg-sorbent/hr

Limestone silo

Limestone convey

Limestone feeder

Limestone blower

To sorbent injection point

Coal

Sorbent silo

Sorbent convey

Sorbent feeder

Sorbent blower

To sorbent injection point

Air

23

Table 9. Mass balance and assumptions of the caliner unit of the SAP.

Item Value Unit

Heat loss 10%

Coal feed rate 1,454 kg-coal/hr

Air coal ratio 0.347 kg-mol-air/kg-coal

Cost methodology and assumptions

Cost is categorized as total project cost (TPC), operating and maintenance cost (O&M),

and equivalent annual cost (EAC). The cost analysis of the DSI unit of the SAP system

was based on the IPM model for a DSI process [12]. The cost analysis of the calciner unit

of the SAP system without an injection system, was based on a pulverized coal-fired

boiler model [13]. Basic parameters and assumptions for the cost calculation are given in

Tables 10 and 11.

Table 10. Cost calculation assumptions.

Variable Label Value Unit Note

Net Output A 550 MWe

Retrofit factor B 1

Limestone feed rate M 17.385 tonne/hr

Limestone cost N 30 $/tonne

Coal cost 52 $/tonne

Sorbent waste rate P 16.8 tonne/hr Based on the final products of

CaSO3, CaSO4, and CaCl2

Aux power Q 0.63 % M*20/A

Fly ash waste rate R 15.4 tonne/hr

Waste disposal cost S 30 $/tonne

Aux power cost T 0.06 $/kWh

Operating labor rate U 60 $/hr Labor costs including all

benefits

Effluent temperature 400 °C Including both sorbent and flue

gas exiting the furnace

TPC cost of the DSI unit of the SAP is a sum of the capital, engineering, and construction

cost (CECC) as well as financing expenditures. Details of these items are presented in

Table 11.

24

Table 11. Total project cost of the DSI unit of the SAP.

Item Label Value Unit Note

BM 16,911,927 $ 7,516,000*(B)*(M)0.284

BM 31 $/kW

Engineering and construction

management costs

A1 845,596 $ 5% of BM

Labor adjustment for 6 x 10

hour shift etc.

A2 845,596 $ 5% of BM

Contractor profit and fees A3 845,596 $ 5% of BM

CECC 19,448,716 $ BM+A1+A2+A3

CECC 35.4 $/kW

Owner’s home office costs

(owner’s engineering,

management, and procurement)

B1 972,436 $ 5% of CECC

Allowance for funds used during

construction

B2 0 $ 0% of the CECC as the

projects are to be

completed <1 year

TPC 20,421,152 $ CECC+B1+B2

TPC 37.1 $/kW

O&M cost of the DSI unit of the SAP includes the fixed (FO&M) and variable (VO&M)

components. The FO&M is further subdivided into the costs associated with additional

operations (FO&MO), maintenance labor and materials (FO&MM), and administrative

and support labor (FO&MA). The VO&M is composed of the reagent cost (VO&MR),

which is a function of the sorbent capacity and acid gas concentrations, waste disposal

cost (VO&MW), and additional auxiliary power required (VO&MP). The disposal cost

accounts for both the sorbent and the fly ash waste. Summary of the O&M calculations is

given in Table 12.

Table 12. O&M cost of the DSI unit of the SAP.

Variable Value Unit Note

FO&MO 0.227 $/(kW yr) (1 additional operator)*2080*U/(A*1000)

FO&MM 0.307 $/(kW yr) BM*0.01/(A*B*1000)

FO&MA 0.010 $/(kW yr) 0.03*(FO&MO+0.4*FO&MM)

Total FO&M 0.54 $/(kW yr) FO&MO+FO&MM+FO&MA

Total FO&M 299,693 $/yr

VO&MR 0.49 $/MWh M*R/A

VO&MW 2.55 $/MWh (N+P)*S/A

VO&MP 0.38 $/MWh Q*T*10

Total VO&M 3.42 $/MWh VO&MR+VO&MW+VO&MP

Total VO&M 16,478,358 $/yr

Total O&M 16,778,050 $/yr

25

TPC and O&M costs of the limestone calciner unit of the SAP are summarized in Table

13. The total O&M cost is overestimated because the SAP unit does not have water and

steam tubes that are included in the boiler model.

Table 13. Cost evaluation for the limestone calciner unit of the SAP.

Item Label Value Unit Note

Equipment cost C1 1,672,679 $ 1979 dollars

Direct installation cost C2 664,390 $ 1979 dollars

indirect installation cost C3 550,217 $ 1979 dollars

TPC 2,887,286 $ C1+C2+C3

TPC 8,672,005 $ 2010 dollars

TPC 15.8 $/kW

Total VO&M D1 100,698 $/yr 1979 dollars, excluding

coal cost

Total FO&M D2 539,087 $/yr 1979 dollars

Total O&M 639,785 $/yr D1+D2, excluding coal

cost

Total O&M D0 1,921,603 $/yr 2010 dollars, excluding

coal

Coal E0 662,157 $/yr

Total O&M 2,583,760 $/yr D0+E0, 2010 dollars

The total project and O&M cost of separate units of the SAP is summarized in Table 14.

The TPC and O&M costs of the DSI unit are 235% and 649% more than their

counterparts of the calciner unit. A much higher cost of O&M in the DSI unit is

associated with greater amounts of limestone handling and processing than that of coal in

the calciner unit (17,385 vs. 1,454 kg/hr).

Table 14. Cost summary of the SAP.

TPC, $ TPC, $/kW Total O&M, $/yr

DSI unit 20,421,152 37.1 16,778,050

Calciner unit 8,672,005 15.8 2,583,760

SAP (DSI and calciner units) 29,093,156 52.9 19,361,810

Cost comparison between the SAP and DSI

The cost of a DSI system applied to Case 9 using commercial quicklime was analyzed

using the same assumptions and methodology for the DSI unit of the SAP discussed

above. The EACs for SAP and DSI systems at different capital cost and equipment life

span scenarios are presented in Table 15. In both the SAP and DSI cases, O&M costs are

5 to 20 times higher than the annualized project costs. Assuming the same sorbent

utilization and acid gas removal rates for either a SAP or a conventional DSI unit, the

26

EACs of the SAP are 36% to 38% lower than the conventional DSI process in the three

scenarios considered in this study. The primarily reason for the lower cost of the SAP is

attributed to cost of producing a lime sorbent in SAP than purchasing it separately for use

with the conventional DSI ($105 vs. $30). In a commercial limestone calcination process,

quicklime is prepared in a rotary kiln operating at 1100 °C and 2-4 hours calcination time.

In contrast, a SAP sorbent is prepared in 1-3 second heat treatment time. These results

demonstrate that SAP is cost-effective technology for acid gas mitigation.

Table 15. Comparison of the costs using the SAP and DSI for the acid gas removal.

O&M, $/yr

Project cost

TPC, $ Interest

of capital

Lifetime of

equipment, yr

Annualized

project cost, $/yr

EAC, $/yr

SAP 19,361,810 29,093,156 5% 15 2,802,901 22,164,712

DSI 34,023,683 17,320,707 5% 15 1,668,717 35,692,399

SAP 19,361,810 29,093,156 10% 15 3,824,987 23,186,798

DSI 34,023,683 17,320,707 10% 15 2,277,219 36,300,902

SAP 19,361,810 29,093,156 10% 20 3,417,271 22,779,082

DSI 34,023,683 17,320,707 10% 20 2,034,484 36,058,167

A sensitivity analysis was also performed to determine the impacts of the sorbent

utilization and price of limestone on the O&M cost of the SAP (Figure 27). The total

O&M cost initially decreases with the increasing calcium utilization of the sorbent and

gradually levels off at about 50% calcium utilization, assuming the same acid gas

removal rates in Table 8. On the other hand, the total O&M increases linearly with the

cost of limestone. Since the transportation fee accounts for a large fraction of the cost of

the limestone, the results indicate that the cost of SAP sorbents is sensitive to the distance

between a quarry and a utility site.

Figure 27. Sensitivity of the total O&M cost to the sorbent utilization and limestone price.

15.0

20.0

25.0

30.0

0% 20% 40% 60% 80%

Tota

l O&

M, M

illio

n $

Sorbent utilization (molar, %)

10.0

15.0

20.0

25.0

30.0

0 20 40 60 80

Tota

l O&

M, M

illio

n $

Limestone Price ($)

27

CONCLUSIONS AND RECOMMENDATIONS

Conclusions

The ISGS prepared lime sorbents showed better capacity for HCl capture, more

than twice of that by commercial Trona and quicklime but exhibited worse

performance for SO3 and SO2 removal.

Partially calcined limestone had the maximum capacity for HCl removal at 70%

extent of calcination (EOC).

The sorbents doped with calcium acetate hydrate [Ca(OAc)2·H2O] had more than

30% higher capacity for HCl capture.

Combination of various sorbents may optimize HCl/SO3/SO2 removal.

Batch sorbents provided useful guidance for producing desired sorbents by

sorbent activation process (SAP).

The SAP technology is cost-effective for on-site production of lime sorbents for

HCl and SO3 removal. SAP’s overall cost of acid gas removal is about 37% lower

than a conventional dry sorbent injection (DSI) process.

Recommendations for Future Work

Recommended future work incudes preparing 100 grams of two sorbents identified in this

project using the SAP and testing the sorbents with an entrained duct reactor with

combustion gases in which the sorbent particles have a residence time similar to that in

power plant applications. This will help in comparing to the results obtained using the

TGA and tubular reactor, and generate more realistic data to advance the SAP technology

towards slipstream testing and eventual commercialization of the technology for acid gas

capture applications.

REFERENCES

1. U.S.EPA, 40 CFR Parts 60 and 63, National Emission Standards for Hazardous

Air Pollutants. February 16, 2012.

2. Ruch, R., H. Gluskoter, and N. Shimp, Occurrence and Distrubution of

Potentially Volatile Trace Elements in Coal, in Enviromental Geology Notes.

August 1974.

3. Presto, A.A. and E.J. Granite, Impact of Sulfur Oxides on Mercury Capture by

Activated Carbon. Environmental Science & Technology, 2007. 41(18): p. 6579-

6584.

4. Partanen, J., et al., Absorption of HCl by Limestone in Hot Flue Gases. Part I: the

Effects of Temperature, Gas Atmosphere and Absorbent Quality. Fuel, 2005.

84(12–13): p. 1664-1673.

5. Chin, T., et al., Hydrated Lime Reaction with HCl under Simulated Flue Gas

Conditions. Industrial & Engineering Chemistry Research, 2005. 44(10): p. 3742-

3748.

6. Yan, R., et al., Kinetic Study of Hydrated Lime Reaction with HCl. Environmental

Science & Technology, 2003. 37(11): p. 2556-2562.

28

7. Abanades, J.C. and D. Alvarez, Conversion Limits in the Reaction of CO2 with

Lime. Energy & Fuels, 2003. 17(2): p. 308-315.

8. Blamey, J., et al., Reactivation of CaO-Based Sorbents for CO2 Capture:

Mechanism for the Carbonation of Ca(OH)2. Industrial & Engineering Chemistry

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9. Lu, Y., Evaluation of Dry Sorbent Technology for Pre-combustion CO2 Capture

(phase II). 2013, Illinois Clean Coal Institute.

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29

DISCLAIMER STATEMENT

This report was prepared by Hong Lu of the Illinois State Geological Survey at the

University of Illinois at Urbana-Champaign (ISGS-UIUC), with support, in part, by

grants made possible by the Illinois Department of Commerce and Economic

Opportunity through the Office of Coal Development and the Illinois Clean Coal

Institute. Neither Hong Lu, ISGS-UIUC, nor any of its subcontractors, nor the Illinois

Department of Commerce and Economic Opportunity, Office of Coal Development, the

Illinois Clean Coal Institute, nor any person acting on behalf of either:

(A) Makes any warranty of representation, express or implied, with respect to the

accuracy, completeness, or usefulness of the information contained in this report,

or that the use of any information, apparatus, method, or process disclosed in this

report may not infringe privately-owned rights; or

(B) Assumes any liabilities with respect to the use of, or for damages resulting from

the use of, any information, apparatus, method or process disclosed in this report.

Reference herein to any specific commercial product, process, or service by trade name,

trademark, manufacturer, or otherwise, does not necessarily constitute or imply its

endorsement, recommendation, or favoring; nor do the views and opinions of authors

expressed herein necessarily state or reflect those of the Illinois Department of

Commerce and Economic Opportunity, Office of Coal Development, or the Illinois Clean

Coal Institute.

Notice to Journalists and Publishers: If you borrow information from any part of this

report, you must include a statement about the state of Illinois’ support of the project.