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Acid Gas Enrichment – More New Wrinkles on an Old Process

CRU International Sulphur Conference 2018 Gothenburg, Sweden

Angie Slavens and Abdulkader Alkasem

UniverSUL Consulting PO Box 109760 Abu Dhabi, UAE

Nathan A. Hatcher1, G. Simon A. Weiland1, and Ralph H. Weiland2

Optimized Gas Treating, Inc. 119 Cimarron Park Loop, Suite A, Buda, TX 78610

12337 Jones Road, Suite 432

Houston, TX 77070

ABSTRACT

Sour gases containing low concentrations of H2S can be particularly troublesome and costly to process in conventional sulphur recovery units (SRUs). Unlike refineries, which routinely operate with 70-90% or higher H2S in the acid gas feeds, sulphur plants in natural gas service often see much lower H2S concentrations. A myriad of problems have been reported through the years with lower H2S feeds, some of which include flame instability and the associated susceptibility of the SRU to contaminants (NH3, hydrocarbons, BTEX) which can lead to operability and reliability concerns, increased operating cost (feed preheat, oxygen enrichment, natural gas co-firing), and difficulty in design. The technique of “enriching” the acid gas, or acid gas enrichment (AGE), overcomes some of these problems, usually at the expense of additional operating equipment and complexity. There are innovative plant line-ups which may leverage AGE benefits, while also optimizing capital expenditure and operating costs. These unique AGE configurations can be utilized to achieve economical, robust operation in grassroots facilities, or can be employed in existing facilities that are required to meet new, lower SO2 regulations. In a previous paper presented at the CRU Middle East Sulphur 2018 Conference, the authors briefly reviewed common AGE configurations and examined several unique AGE line-ups from a techno-economic perspective. By understanding the basic operating parameters that affect selectivity through mass transfer rate-based simulation, several selectivity improvement opportunities that have been previously unpublished were evaluated. In this paper, the authors expand upon the previous work by evaluating the alternative configurations at leaner acid gas feed concentrations. Comparisons to other conventional schemes such as natural gas co-firing and O2-enrichment are also investigated.

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INTRODUCTION

Acid gas enrichment (AGE) has been applied successfully for many years to upgrade the quality of sour gases containing low concentrations of H2S, which can be particularly troublesome and costly to process in conventional sulphur recovery units (SRUs). Lower H2S concentration feeds can lead to flame instability in the SRU front-end main burner and increase susceptibility of the SRU to contaminants (NH3, hydrocarbons, BTEX), which can lead to operability and reliability concerns. In simplistic terms, AGE addresses these problems by employment of an amine unit (usually MDEA-based) that is selective for H2S, to produce an acid gas that is upgraded, or richer in H2S, to be processed in the sulphur plant. While acid gas enrichment provides many technical advantages, its capital and operating costs can be challenging to overcome. Various techniques have been employed to optimize the flowsheet configuration to minimize lifecycle cost, and this paper explores several unique alternatives. Besides acid gas enrichment, other measures can be taken to boost reaction furnace temperature, including feed pre-heat, natural gas co-firing, and oxygen enrichment. These measures also increase capital and operating costs. Additionally, an SRU processing lean acid gas must carry with it the extra baggage of CO2 that comes along for a ride, creating operational challenges that are discussed throughout this paper. The issue of lean acid gases is almost entirely constrained to gas plants. Refineries routinely process 70-90% H2S and do not usually face reaction furnace temperature or flame stability problems. BACKGROUND The primary aim of acid gas enrichment is to maximize H2S content in the enriched acid gas stream that flows to the SRU. The level of achievable enrichment is dependent upon several factors, including partial pressure of H2S in the lean acid gas feeding the enrichment absorber and selectivity of the solvent. The configurations explored in this paper accomplish increased H2S partial pressure via the following methods:

Mixing lean acid gas with SRU tail gas to produce a combined stream with higher H2S partial pressure than a typical tail gas treating (TGT) absorber feed stream

Mixing enriched acid gas (from the regenerator) with lean acid gas to produce a stream with higher H2S partial pressure than a typical AGE or TGT absorber feed stream (HIGHSULF PLUSTM configuration, T. K. Kanmamedov and R. H. Weiland, 2013)

Leaving more sulphur in the SRU tail gas, increasing H2S partial pressure in the TGT absorber feed stream

The original incarnation of this study explored various configurations of the above using a lean acid gas stream containing 30 mol% H2S (dry basis). After presenting the results at Middle East Sulphur 2018, it became evident that it would also be beneficial to explore whether the alternative schemes could provide greater benefits when processing a leaner acid gas feed (15 mol% H2S , dry basis). In addition, audience questions after the presentation prompted the authors to compare the results of the AGE cases to other conventional methods such as co-firing and O2 enrichment.

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DESIGN BASIS AND PROCESS CONFIGURATIONS CONSIDERED

A 1,000 MTPD sulphur recovery unit, with nominal 30 mol% H2S (dry basis) feed composition, typical of a sour gas plant, is considered as the basis for comparing several process design configurations. The wet basis composition of this stream is provided in Table 1. Table 1. Lean Acid Gas (LAG) Feed Stream Design Basis

Component Unit Lean Acid Gas

(LAG)

Water mole % 10.476

Hydrogen Sulfide mole % 26.855

Carbon Dioxide mole % 62.183

Methane mole % 0.268

Ethane mole % 0.100

Propane mole % 0.069

n-Butane mole % 0.029

Isobutane mole % 0.009

n-Pentane mole % 0.002

Isopentane mole % 0.003

n-Hexane mole % 0.003

Total kmole/hr 4,838.174

Temperature Celsius 60.0

Pressure barg 0.92

Table 2 lists key design features of the process configurations evaluated, each of which is illustrated and further described in Figures 1 through 7 that follow.

Table 2. Process Configuration Cases

Case Basic Description AGE Absorber LAG Feed To Enriched AG

(EAG) Feed To Claus

Stages

Base Conventional AGE Standalone AGE Absorber SRU RF 2

Case A Common AGE/TGT Absorber Combined AGE/TGT

AGE/TGT Absorber

SRU RF 2

Case B(1) Common Absorber + EAG Recycle

Combined AGE/TGT

Quench Column

50% SRU RF / 50% Quench

2

Case C Common Absorber + Single Claus Stage SRU

Combined AGE/TGT

AGE/TGT Absorber

SRU RF 1

Case D Case C with LAG Feed to Quench Column

Combined AGE/TGT

Quench Column

SRU RF 1

Case E Co-firing + Single Claus Stage SRU

None SRU RF SRU RF 1

Case F O2-enrichment + Single Claus Stage SRU

None SRU RF SRU RF 1

(1) HIGHSULF PLUSTM configuration (T. K. Kanmamedov and R. H. Weiland, 2013)

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A 40 wt% MDEA solvent solution, containing 0.4 wt% phosphoric acid, is considered for AGE and TGT solvent in all cases. Hot ambient conditions, representative of the Middle East, are considered. Thus, lean amine cooling is accomplished in two stages - bulk air cooling to 55oC followed by trim cooling, with propane refrigeration, to 45oC. The conventional acid gas enrichment Base Case process configuration is shown in Figure 1. Acid gas from the upstream sour gas treatment unit is first processed in an AGE absorber, which absorbs H2S from the acid gas using MDEA. AGE absorber overhead gas, containing most of the CO2 in the acid gas and other components not absorbed by the amine (including hydrocarbons), flows directly to the tail gas incinerator. Rich amine from the separate AGE and TGT absorbers is routed to the common regenerator, which produces concentrated acid gas, containing all of the sulfur to be recovered, plus the tail gas unit recycled sulphur. This enriched acid gas stream flows to the SRU/TGTU for sulphur recovery. The SRU is equipped with two Claus stages, as in a conventional design.

Figure 1. Base Case (Conventional AGE) PFD

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The Case A process configuration is shown in Figure 2. This scheme is identical to the Base Case but achieves higher H2S partial pressure in the TGT absorber by mixing acid gas and SRU tail gas prior to absorption in a single absorber. In addition to improving enrichment, this scheme reduces equipment count/capital cost by combining the enrichment and tail gas treating absorbers in a single, but larger, absorber vessel. The combined AGE/TGT absorber overhead gas flows directly to the incinerator and rich amine from the common absorber is routed to the common regenerator. By floating the AGE/TGT Absorber onto the Incinerator, operating pressure in the combined absorber is minimized, which works to maximize selectivity. Similar to the Base Case, the regenerator produces concentrated acid gas, containing all of the sulfur to be recovered plus the tail gas unit recycled sulphur. This enriched acid gas stream flows to the SRU/TGTU for sulphur recovery. The SRU is equipped with two Claus stages, as in a conventional design. Figure 2. Case A PFD

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The Case B process configuration is shown in Figure 3. This scheme is similar to Case A but recycles 50% of the enriched acid gas back through the combined AGE/TGT absorber, in order to further increase H2S concentration in the enriched acid gas stream (HIGHSULF PLUSTM, T. K. Kanmamedov and R. H. Weiland, 2013). In addition, the lean acid gas stream is routed to the quench column for cooling and partial water condensation/removal to improve absorption efficiency. The combined AGE/TGT absorber overhead gas flows directly to the incinerator and rich amine from the absorber is routed to a common regenerator. The regenerator produces concentrated acid gas, containing all of the sulfur to be recovered plus the tail gas unit recycled sulphur. Half of this enriched acid gas stream flows to the SRU/TGTU for sulphur recovery, while the other half is recycled back to the absorber for further enrichment, as described above. The acid gas recycle ratio is an optimization parameter; the 50% recycle basis is taken simply for illustrative purposes. The SRU is equipped with two Claus stages, as in a conventional design. Figure 3. Case B PFD

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The Case C process configuration is shown in Figure 4. This scheme is identical to Case A, but the second Claus stage has been eliminated in order to increase SRU tail gas H2S concentration. The richer tail gas increases H2S partial pressure in the AGE/TGT absorber, resulting in a higher H2S concentration in the enriched acid gas. Removal of a Claus stage also reduces capital cost. It should be noted that this change reduces SRU sulphur recovery efficiency, which may be a concern during TGTU bypass operation. This will be discussed further in the process analysis section of the paper. In Case C, the final Claus condenser is a low pressure (LP) steam generator, rather than a low low pressure (LLP) steam generator, to maximize sulphur in the tail gas, as well as to maximize the recovery of waste heat as valuable LP steam rather than rejecting to the atmosphere.

Figure 4. Case C PFD

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The Case D process configuration is shown in Figure 5. This scheme is similar to Case B but with only a single Claus stage in the SRU. Additionally, all of the enriched acid gas is recycled to the SRU and no enriched acid gas is combined with the tail gas (HIGHSULF PLUSTM configuration eliminated). Similar to Case C, the final Claus condenser is an LP steam generator, rather than an LLP steam generator, to maximize sulphur in the tail gas and recover waste heat as useful LP steam.

Figure 5. Case D PFD

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The Case E process configuration is shown in Figure 6. This scheme does not employ AGE, but rather considers a conventional SRU/TGTU with natural gas co-firing to achieve reaction furnace temperatures similar to that which can be achieved with acid gas enrichment. Due to increased COS formation in the reaction furnace, a COS hydrolysis reactor is required in the TGTU to achieve the target sulphur recovery efficiency (SRE) of 99.9%. For the purpose of consistent comparison, only a single Claus stage is employed in the SRU; however, this is not likely to be realistic for TGTU bypass operation because Claus recovery efficiency is expected to be significantly lower than in the AGE cases due to higher CO2 concentration and increased COS formation. This will be discussed further in the process analysis section of the paper. Figure 6. Case E PFD

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The Case F process configuration is shown in Figure 7. This scheme considers a conventional SRU/TGTU employing 100% O2-enrichment. For the Lean Acid Gas feed stream, even with 100% O2 enrichment and maximum acid gas preheating, it is not possible to achieve reaction furnace temperatures similar to the AGE cases; thus, supplemental natural gas co-firing is also required. Due to increased COS formation in the reaction furnace, a COS hydrolysis reactor is required in the TGTU to achieve the target sulphur recovery efficiency (SRE) of 99.9%. For the purpose of consistent comparison, only a single Claus stage is employed in the SRU; however, this is not likely to be realistic for TGTU bypass operation because Claus recovery efficiency is expected to be significantly lower than in the AGE cases due to higher CO2 concentration and increased COS formation. This will be discussed further in the process analysis section of the paper.

Figure 7. Case F PFD

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PROCESS PARAMETERS ANALYSIS ProTreat version 6.4.2.003 was used to simulate the cases described above. Table 3 provides a summary of key process parameters, considering the Lean Acid Gas (nominal 30 mol% H2S) feed stream. Table 3. Key Process Performance Parameters for 30 mol% H2S Lean Acid Gas Feed

Process Parameter Base Case A Case B Case C Case D Case E Case F

Lean Amine to AGE Absorber (m3/hr)

1,164 --- --- --- --- --- ---

Lean Amine to TGT Absorber (m3/hr)

454 --- --- --- --- --- ---

Total Lean Amine from Regen. (m3/hr)

1,709 2,095 3,087 2,155 2,131 1,666 879

H2S in Absorber Inlet (mol%)

26.8% (AGE) 3.2% (TGT)

17.2% 28.9% 17.6% 17.6% 1.8% (TGT) 4.5% (TGT)

H2S Removed in Absorber (kmol/hr)

1,298 (AGE) 115 (TGT)

1,415 2,830 1,454 1,450 267 (TGT) 240 (TGT)

H2S in Enriched AG (mol%)

67.0% 74.3% 80.2% 74.3% 74.4% 46.4% 43.6%

Natural Gas Co-firing (kmol/hr)

--- --- --- --- --- 871 272

O2 (kmol/hr) --- --- --- --- --- --- 816

Reaction Furnace Temperature (oC)

1,093 1,148 1,192 1,142 1,143 1,154 1,128

Claus SRU Tail Gas Flow (kmol/hr)

4,487 4,276 4,185 4,408 4,390 15,651 6,332

COS in SRU Tail Gas (kmol/hr)

5.4 4.2 2.9 4.7 4.3 35.8 32.7

Claus SRE (EOR) 91.8% 91.7% 91.8% 89.3% 89.5% 82.9% 84.3%

Overall SRE (EOR) 99.9% 99.9% 99.9% 99.9% 99.9% 99.9% 99.9%

Amine Circulation The amine circulation data provided in Table 3 are plotted in Figure 8 and illustrate that the combined AGE/TGT absorber cases require higher amine circulation rates than the Base Case. This is due to the fact that mixing the tail gas with the acid gas feed dilutes the acid gas feed stream from 26.8 mol% H2S to approximately 17 mol% H2S. Additionally, the pressure of the AGE absorber is reduced from approximately 0.9 to 0.3 barg in the combined absorber cases. Both of these factors contribute to lower H2S partial pressure in the AGE absorber, which increases lean amine circulation requirements. On the other hand, the H2S partial pressure of the combined stream (Cases A-D) is higher than that which would be achieved in a standalone TGT absorber (approximately 17 versus 3 mol%), which decreases TGT absorber amine circulation requirements. However, since the quantity of H2S removed

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in the AGE absorber is more than ten-fold the quantity to be removed in the TGT absorber, the net impact is an overall increase in circulation rate. The circulation rate for Case B is higher than all other cases, due to the additional H2S input from the enriched acid gas stream. Recycling 50% of the enriched acid gas increases the feed stream H2S content to approximately 29 mol%, but also doubles the quantity of H2S to be removed. The circulation rates for Cases E and F are lower than for the enrichment cases. However, the amine circulation rate of the Case E TGT absorber is nearly four times greater than that required for the Base Case TGT absorber. This is due to much lower H2S partial pressure resulting from dilution by natural gas co-firing. By utilizing O2 in Case F, H2S dilution is reduced significantly; however, its circulation rate is still double the Base Case value due to a leaner SRU feed stream and supplemental natural gas co-firing. Figure 8. Lean Amine Circulation Comparison for Lean Acid Gas Feed (30 mol% H2S)

Acid Gas Enrichment Level, Reaction Furnace Temperature and SRU Size Absorber inlet and enriched acid gas H2S concentrations for all cases are plotted in Figure 9, which illustrates that as absorber inlet H2S concentration increases, so too does the level of acid gas enrichment.

1,709

2,095

3,087

2,155 2,131

1,666

879

0

500

1,000

1,500

2,000

2,500

3,000

3,500

Base Case Case A Case B Case C Case D Case E Case F

Tota

l Lea

n A

min

e C

ircu

lati

on

(m3/h

)

13

Figure 9. Level of Enrichment Achieved for Lean Acid Gas Feed (30 mol% H2S)

The degree of acid gas enrichment determines reaction furnace temperature, as illustrated in Table 3. A minimum temperature of around 1,050oC is considered for hydrocarbon destruction, although it could be argued that temperatures as low as 1,000oC could be acceptable in the enrichment cases, as hydrocarbons will slip through the absorber and flow straight to the incinerator. This is not the case for Cases E and F, where all acid gas feed contaminants flow through the reaction furnace and must be destroyed. For this reason, these cases require supplemental firing, with natural gas and/or oxygen. The degree of enrichment also impacts the size of the SRU/TGTU. The greater the H2S concentration, the less CO2 flow through the unit, which decreases equipment/piping size, as illustrated by the tail gas flow rates in Table 3. The tail gas flow rate of Case E is more than three times greater than the enrichment cases due to the combined impact of leaner acid gas feed and natural gas co-firing. Utilization of O2 enrichment in Case E mitigates the impact of leaner acid gas by removing nitrogen; however, CO2 in the acid gas still flows through the unit. This, combined with supplemental natural gas co-firing, leads to a tail gas flow rate that is approximately 50% greater than the enrichment cases. Sulphur Recovery Efficiency As illustrated in Table 3, end of run (EOR) sulphur recovery efficiency in the Claus SRU is approximately 92% for the Base Case and Cases A and B. Removing one of the two Claus stages (Cases C and D) reduces SRE by approximately 2.5%, depending on the operating conditions of the SRU. In Cases E and F, the removal of the second Claus stage reduces SRE by an additional 5-6% due to a combination of leaner acid gas and greater COS formation in the reaction furnace. Elimination of the second Claus stage may cause concerns related to high SO2 emissions during TGTU bypass mode; however, the impact will be significantly more pronounced in the non-AGE cases (Cases E and F). Thus, while only a single Claus stage has been considered for Cases E and F for the sake of comparison consistency, this is not likely to be a feasible arrangement.

67.0

74.3

80.2

74.3 74.4

46.643.6

26.9

17.2

28.9

17.6 17.6

1.84.5

0

10

20

30

40

50

60

70

80

90

Base Case Case A Case B Case C Case D Case E Case F

H2S

Co

nce

ntr

atio

n (m

ol%

)

(AGE)

3.2(TGT)

- H2S in Lean AG Feed to Absorber

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COMMERCIAL COMPARISON

Table 4 provides a summary of key commercial parameters for the cases studied, considering the Lean Acid Gas (30 mol% H2S) feed stream. Table 4. Key Commercial Parameters for Lean Acid Gas Feed Stream (30 mol% H2S)

Parameter Base Case A Case B Case C Case D Case E Case F

CapEx ($Million USD)

AGEU 20.8 28.6 31.8 28.6 28.6 --- ---

SRU (Case F includes ASU) 104 102 100 91.6 91.4 190 150

TGTU (Cases E/F include COS Rx)

107 105 121 106 106 172 106

Total CapEx 232 235 253 226 226 362 256

OpEx ($Million USD per Annum)

HP Steam ($5.8/ton) (8.6) (8.8) (9.0) (9.2) (9.2) (20.9) (8.7)

LP Steam ($3.6/ton) 5.9 7.5 11.6 7.7 7.6 3.0 1.3

Fuel Gas ($2.6/MMBtu) 5.4 5.5 5.6 5.6 5.6 22.8 6.8

Power ($70/MWh) 8.5 9.6 13.9 9.8 9.9 13.5 6.2

Oxygen ($50/ton) --- --- --- --- --- --- 11.1

Catalyst & Chemicals 1.2 1.2 1.4 1.2 1.2 4.0 1.7

Maintenance (5% CapEx * # equip. factor)

11.6 10.9 11.8 9.3 9.3 14.8 10.5

Total OpEx 23.8 25.9 35.3 24.5 24.4 37.3 29.0

Lifecycle Cost, 30 years @ 6% ($Million USD)

Total LCC 560 591 739 563 561 875 654

Capital Cost The capital cost data provided in Table 4 are plotted in Figure 10, and illustrate that the CapEx of the enrichment cases are all fairly similar, with the exception of Case B which is costlier due to the requirement to remove double the H2S quantity in the AGE/TGT absorber. The smaller Claus unit equipment with the richer acid gas feed from Case B is not enough to offset the higher amine unit cost. Despite the fact that Case E does not require an AGE absorber and thus has lower amine circulation requirements, its much larger process gas volumetric flow increases SRU/TGTU size and capital cost. This case also has additional COS hydrolysis reactor costs. The capital cost of Case F is lower than Case E; however, the air separation unit (ASU) required for O2 production, and COS hydrolysis reactor, largely negate its cost advantages.

15

16

Figure 10. Capital Cost Comparison for Lean Acid Gas Feed (30 mol% H2S)

Operating Cost The operating cost data provided in Table 4 are plotted in Figure 11, and illustrate that the OpEx of the enrichment cases are all fairly similar, with the exception of Case B, which has higher LP steam and power costs due to its higher circulation rate. Fuel gas and power costs for Case E are significant due to co-firing, and catalyst costs are higher due to greater volumetric flow and the additional COS hydrolysis reactor. These higher operating costs are only partially offset by additional HP steam production (due to co-firing) and lower LP steam consumption (due to no AGE). Case F benefits from lower LP steam and power costs, due to low circulation rate, but these benefits are offset by oxygen production costs. Figure 11. Operating Cost Comparison for Lean Acid Gas Feed (30 mol% H2S)

231 235253

226 226

362

256

0

50

100

150

200

250

300

350

Base Case Case A Case B Case C Case D Case E Case F

Cap

Ex (m

illi

on

USD

)

23.825.9

35.332.8

24.4

37.3

28.9

0

5

10

15

20

25

30

35

40

Base Case Case A Case B Case C Case D Case E Case F

Op

Ex (m

illi

on

USD

per

an

nu

m)

17

Lifecycle Cost The lifecycle cost data provided in Table 4 are plotted in Figure 12, and illustrate that the lifecycle costs of Cases C and D are very similar to the Base Case. Case D is slightly more attractive than Case C due to acid gas cooling and water condensation in the quench column. As expected, lifecycle cost for Case E is higher than all other cases. Lifecycle cost for the O2 enrichment case (Case F) is higher than the acid gas enrichment cases; however, it must be noted that this case requires both O2 enrichment and natural gas co-firing due to insufficient furnace temperature with O2 alone. It could be interesting to consider a case that combines mild acid gas enrichment and O2 enrichment to avoid co-firing. Figure 12. Lifecycle Cost Comparison for Lean Acid Gas Feed (30 mol% H2S)

Given the similar lifecycle costs of the Base Case and Case D, one would rightly question why the alternative configuration would be selected over the Base Case. Enrichment is improved and reaction furnace temperature is higher for Case D, but this does not lead to significant technical or commercial advantages. To the contrary, a less attractive feature of Case D is that it will result in higher SO2 emissions during TGTU bypass mode due to the presence of only a single Claus stage in the SRU. Thus, at this point in the analysis, the authors believed it would be prudent to investigate whether Case D may offer benefits at leaner acid gas concentrations.

328357

485

337 335

514398

0

100

200

300

400

500

600

700

800

900

1,000

Base Case Case A Case B Case C Case D Case E Case F

Life

cycl

e C

ost

(mill

ion

USD

)

CapExCapEx

CapEx

CapEx

CapEx

CapEx

CapEx

OpExOpEx

OpEx

OpEx OpEx

OpExOpEx

560591

739

563 561

875

654

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IMPACT OF LEANER ACID GAS To determine whether Case D may have benefits when processing acid gas with H2S concentration below 30 mol%, an Ultra-Lean Acid Gas case (nominal 15 mol% H2S, dry basis) was studied. This represents the lowest viable concentration to maintain acceptable reaction furnace temperature when using MDEA for acid gas enrichment. Wet basis composition for the Ultra-Lean Acid Gas stream is provided in Table 5 and resulting key process parameters are summarized in Table 6. Table 5. Ultra-Lean Acid Gas Feed Stream Design Basis

Component Unit Ultra-Lean Acid Gas

Water mole % 9.479

Hydrogen Sulfide mole % 13.579

Carbon Dioxide mole % 76.452

Methane mole % 0.271

Ethane mole % 0.102

Propane mole % 0.070

n-Butane mole % 0.030

Isobutane mole % 0.009

n-Pentane mole % 0.002

Isopentane mole % 0.003

n-Hexane mole % 0.003

Total kmole/hr 9,570.261

Temperature Celsius 60.0

Pressure barg 0.92

Table 6. Key Process Performance Parameters for 15 mol% H2S Ultra-Lean Acid Gas Feed

Process Parameter Base Case Ultra-Lean

Case D Ultra-Lean

Lean Amine to AGE Absorber (m3/hr) 1,828 ---

Lean Amine to TGT Absorber (m3/hr) 795 ---

Total Lean Amine from Regen. (m3/hr) 2,770 3,029

H2S in Absorber Inlet (mol%) 13.6% (AGE) 2.9% (TGT)

10.7%

H2S Removed in Absorber (kmol/hr) 1,299 (AGE) 125 (TGT)

1,442

H2S in Enriched AG (mol%) 52.5% 61.0%

Reaction Furnace Temperature (oC) 1,034 1,092

Claus SRU Tail Gas Flow (kmol/hr) 5,051 4,713

COS in SRU Tail Gas (kmol/hr) 10.1 7.5

Claus SRE (EOR) 91.2% 90.0%

Overall SRE (EOR) 99.9% 99.9%

Commented [sw1]: Is text meant to be small?

19

As illustrated in Figure 13, amine circulation differential between the Base Case and Case D has decreased significantly for the Ultra-Lean Acid Gas case. This is due to the fact that the H2S concentration of the combined stream feeding the AGE/TGT absorber in Case D is now much closer to the H2S concentration in the feed to the conventional AGE absorber in the Base Case. In other words, the AGE dilution effect is less significant and amine circulation rate is less negatively impacted. Figure 13. Amine Circulation Comparison for Ultra-Lean Acid Gas Feed (15 mol% H2S)

As previously observed, Case D results in improved acid gas enrichment, higher reaction furnace temperature and smaller SRU/TGTU size. The Base Case reaction furnace temperature is marginally acceptable, at below 1,050oC, but is still assumed feasible since most hydrocarbon contaminants will slip through the AGE absorber to the incinerator. Table 7 and Figures 14 through 16 provide a commercial comparison between the Base Case and Case D for the Ultra-Lean Acid Gas feed stream, and reveal the following benefits:

CapEx advantage of Case D improves from 2.6% to 5.8%

OpEx disadvantage of 2.5% for Case D converted to 5.4% advantage

Lifecycle cost slight disadvantage for Case D converted to 5.2% advantage

1,709

2,131

2,7703,029

0

500

1,000

1,500

2,000

2,500

3,000

3,500

4,000

Base Case,Lean

Case D,Lean

Base Case,Ultra-Lean

Case D,Ultra-Lean

Tota

l Lea

n A

min

e C

ircu

lati

on (m

3/h

)

+24.7%

+9.4%

20

Table 7. Key Commercial Parameters for Ultra-Lean Acid Gas Feed Stream

Parameter Base Case Ultra-Lean

Case D Ultra-Lean

CapEx ($Million USD)

AGEU 31.3 38.4

SRU 116 100

TGTU 130 123

Total CapEx 277 261

OpEx ($Million USD per Annum)

HP Steam ($5.8/ton) (13.7) (13.5)

LP Steam ($3.6/ton) 10.1 11.1

Fuel Gas ($2.6/MMBtu) 10.9 10.7

Power ($70/MWh) 13.0 13.2

Catalyst & Chemicals 1.4 1.4

Maintenance (5% CapEx * # equip. factor)

13.8 10.7

Total OpEx 35.5 33.6

Lifecycle Cost, 30 years @ 6% ($Million USD)

Total LCC 766 726

Figure 14. Capital Cost Comparison for Ultra-Lean Acid Gas Feed (15 mol% H2S)

231 226

277262

0

50

100

150

200

250

300

350

Base Case,Lean

Case D,Lean

Base Case,Ultra-Lean

Case D,Ultra-Lean

Cap

Ex (m

illio

n U

SD)

-2.6%

-5.8%

21

Figure 15. Operating Cost Comparison for Ultra-Lean Acid Gas Feed (15 mol% H2S)

Figure 16. Lifecycle Cost Comparison for Ultra-Lean Acid Gas Feed (15 mol% H2S)

Given its lower lifecycle cost, enhanced enrichment and higher reaction furnace temperature, Case D may be more attractive than the Base Case when acid gas leaner than 30 mol% H2S is processed. Additionally, given the smaller differential in SRU recovery efficiency (1.2% reduction versus 2.3% reduction), there is less SO2 emission risk associated with TGTU bypass mode operation for Case D which is equipped with only a single Claus stage.

23.8 24.4

35.533.7

0

5

10

15

20

25

30

35

40

45

Base Case,Lean

Case D,Lean

Base Case,Ultra-Lean

Case D,Ultra-Lean

Op

Ex (m

illio

n U

SD p

er a

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um

)

+2.5%

-5.1%

560 561

766726

0

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Base Case,Lean

Case D,Lean

Base Case,Ultra-Lean

Case D,Ultra-Lean

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+0.2%

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22

UTILITY PRICE SENSITIVITY ANALYSIS Operating costs per utility service are provided in Figure 17, which illustrates that electric power contributes significantly to overall utility cost. The $70/MWh price used in this study is based on typical cost to purchase power from the grid in the Middle East, which is much higher than in other regions of the world. Also, a power price of $70/MWh is inconsistent with a fuel price of $2.6/MMBtu. Thus, an alternative utility cost comparison, with power cost of $25/MWh, is provided in Figure 18 for the purpose of sensitivity analysis. In this case, oxygen price has been reduced from $50/ton to $40/ton to account for the power cost that is included in this figure. Figure 17. Utility Cost Comparison for Lean Acid Gas Feed ($70/MWh Power Price)

Figure 18. Utility Cost Comparison for Lean Acid Gas Feed ($25/MWh Power Price)

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EX (

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Base Case Case A Case B Case C Case D Case E Case F

HP Steam LP Steam Fuel Gas Power Oxygen$5.80/ton $3.60/ton $2.60/MMBtu $70/MWh $50/ton

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HP Steam LP Steam Fuel Gas Power Oxygen$5.80/ton $3.60/ton $2.60/MMBtu $25/MWh $40/ton

23

Resulting lifecycle costs considering the revised power and oxygen prices are provided in Figures 19 and 20. The general trends are the same as previously observed; however, since lifecycle cost is now slightly less affected by OpEx, the attractiveness of Case D has improved slightly. The improvement is thought to be within the margin of error of the calculations, however. Thus, despite the fact that higher than average power price was used in this study, the net impact is negligible and it can be concluded that the comparison is fairly insensitive to power price. Figure 19. Lifecycle Cost Comparison for Lean Acid Gas Feed ($25/MWh Power Price)

Figure 20. Lifecycle Cost Comparison for Ultra-Lean Acid Gas Feed ($25/MWh Power Price)

Finally, an exercise was carried out to determine the impact resulting from operation in a cooler climate, where refrigeration is not required to achieve lean amine temperatures of 45oC. While the

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Base Case Case A Case B Case C Case D Case E Case F

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CapExCapEx CapEx

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OpEx OpEx

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OpExOpEx

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-2.2%

-6.3%

24

results are not tabulated herein, when lower ambient temperature is combined with lower power price ($25/MWh), the lifecycle cost advantage of Case D (versus the Base Case) improves slightly further to 3.3% for Lean Acid Gas and X.X% for Ultra-Lean Acid Gas. CONCLUSIONS When it comes to the challenges associated with processing lean acid gas, acid gas enrichment provides reliability, operability and commercial advantages compared to other conventional schemes such as natural gas co-firing and O2 enrichment. While it is true that O2 enrichment is likely to provide advantages over natural gas co-firing, this feature alone cannot achieve adequate reaction furnace temperature at acid gas H2S concentrations below 30 mol%. Supplemental natural gas co-firing and/or AGE must also be employed, both of which increase operating complexity and cost. Despite its many technical advantages, acid gas enrichment has inherent commercial challenges related to high amine circulation requirements. Various unique processing schemes have been explored in an attempt to overcome these challenges. In all cases, enrichment is improved, but in some cases the resulting amine circulation increase negates commercial benefits. However, when acid gas concentration drops below 30 mol% (dry basis), it may be commercially advantageous to consider a scheme which employs a common AGE/TGT absorber and only a single Claus SRU stage (Cases C and D). The removal of a Claus stage increases SO2 emission risk during TGTU bypass mode but the impact is fairly negligible and could be easily managed by operators who are adept at avoiding this mode of operation, which is becoming more common as governments impose stricter emissions regulations, even for off-design operating conditions. NOMENCLATURE

AG acid gas LLP low low pressure

AGE acid gas enrichment MDEA methyl diethanolamine

AGRU acid gas removal unit mol% mole percent

ASU air separation unit MTPD metric tons per day

barg bar gauge MWh megawatt hours

BFW boiler feed water NH3 ammonia

BTEX benzene, toluene, ethylbenzene, xylene

O2 oxygen

CapEx capital expenditure OpEx operating expenditure

CO carbon monoxide PFD process flow diagram

CO2 carbon dioxide RF reaction furnace

COS carbonyl sulfide RGG reducing gas generator

EAG enriched acid gas SO2 sulphur dioxide

EOR end of run SRE sulphur recovery efficiency

H2 hydrogen SRU sulphur recovery unit

H2S hydrogen sulphide TGT tail gas treating

HP high pressure TGTU tail gas treating unit

kmol kilomole ULAG ultra-lean acid gas

LAG lean acid gas WHB waste heat boiler

LP low pressure

Commented [sw2]: Is this value for OGT to fill out or just left as a place holder?

25

Kg kilogram wt% weight percent

REFERENCES 1. T. K. Khanmamedov and R. H. Weiland, “New AGE Strategies with HIGHSULFTM,” Sulphur Magazine

No. 346, May-June 2013. 2. A. Slavens, S. Weiland, N. Hatcher, A. Alkasem and H. Dhinda, “Acid Gas Enrichment – New

Wrinkles on an Old Process,” presented at CRU Middle East Sulphur 2018, Abu Dhabi, UAE.