JP4573106B2 - A method for separating and recovering ammonia and hydrogen sulfide from desulfurization wastewater. - Google Patents

A method for separating and recovering ammonia and hydrogen sulfide from desulfurization wastewater. Download PDF

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JP4573106B2
JP4573106B2 JP2004367675A JP2004367675A JP4573106B2 JP 4573106 B2 JP4573106 B2 JP 4573106B2 JP 2004367675 A JP2004367675 A JP 2004367675A JP 2004367675 A JP2004367675 A JP 2004367675A JP 4573106 B2 JP4573106 B2 JP 4573106B2
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stripper
ammonia
hydrogen sulfide
ammonia recovery
water
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JP2006169079A (en
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昌弘 小川
祐作 瀧田
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昌弘 小川
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    • C—CHEMISTRY; METALLURGY
    • C02—TREATMENT OF WATER, WASTE WATER, SEWAGE, OR SLUDGE
    • C02F—TREATMENT OF WATER, WASTE WATER, SEWAGE, OR SLUDGE
    • C02F1/00—Treatment of water, waste water, or sewage
    • C02F1/20—Treatment of water, waste water, or sewage by degassing, i.e. liberation of dissolved gases
    • C—CHEMISTRY; METALLURGY
    • C02—TREATMENT OF WATER, WASTE WATER, SEWAGE, OR SLUDGE
    • C02F—TREATMENT OF WATER, WASTE WATER, SEWAGE, OR SLUDGE
    • C02F2101/00—Nature of the contaminant
    • C02F2101/10—Inorganic compounds
    • C02F2101/101—Sulfur compounds
    • C—CHEMISTRY; METALLURGY
    • C02—TREATMENT OF WATER, WASTE WATER, SEWAGE, OR SLUDGE
    • C02F—TREATMENT OF WATER, WASTE WATER, SEWAGE, OR SLUDGE
    • C02F2101/00—Nature of the contaminant
    • C02F2101/10—Inorganic compounds
    • C02F2101/16—Nitrogen compounds, e.g. ammonia
    • C—CHEMISTRY; METALLURGY
    • C02—TREATMENT OF WATER, WASTE WATER, SEWAGE, OR SLUDGE
    • C02F—TREATMENT OF WATER, WASTE WATER, SEWAGE, OR SLUDGE
    • C02F2103/00—Nature of the water, waste water, sewage or sludge to be treated
    • C02F2103/34—Nature of the water, waste water, sewage or sludge to be treated from industrial activities not provided for in groups C02F2103/12 - C02F2103/32
    • C02F2103/36—Nature of the water, waste water, sewage or sludge to be treated from industrial activities not provided for in groups C02F2103/12 - C02F2103/32 from the manufacture of organic compounds
    • C02F2103/365—Nature of the water, waste water, sewage or sludge to be treated from industrial activities not provided for in groups C02F2103/12 - C02F2103/32 from the manufacture of organic compounds from petrochemical industry (e.g. refineries)

Description

本発明は、石油精製装置の重質油脱硫装置から排出した脱硫処理廃水からのアンモニアと硫化水素の分離回収方法に関するものである。 The present invention relates to a method for separating and recovering ammonia and hydrogen sulfide from desulfurization wastewater discharged from a heavy oil desulfurization unit of an oil refinery.

従来、石油精製産業では国内外を問わず原油中の窒素、硫黄はガソリン、灯油軽油などの製品規格を満足するために精製過程で水素化処理され除去される。
その水素化処理過程で窒素及び硫黄は、それぞれアンモニア(NH3)と硫化水素(H2S)に変換される。通常、この硫化水素と副生アンモニアを含むガスは水洗され洗浄水(液相)に吸収され廃水として廃水処理設備( Foul Water Treating Unit )にて処理され廃水からNH3とH2Sの混合ガスが蒸留分離される。
図６に製油所における一般的な公知の廃水処理設備のフローを示す。
この廃水処理設備において、蒸留塔の既存廃水ストリッパー０１の塔頂からでるNH3とH2Sの混合ガスは冷却器０２と還流ドラム０３を介して硫黄回収装置(SRU)０４に送られ、ここで硫化水素は、固形硫黄として製品化される。
アンモニアは、硫黄回収装置(SRU)０４の加熱炉（図示せず）で燃焼して酸化窒素ガスとして大気に放出される一方、処理後排水中に残存し海に放出される。
このため、該酸化窒素ガス放出よる大気汚染の問題、アンモニアの海水放出による海水富栄養化と赤潮発生の問題を惹起していた。
そこで大気汚染防止策及び海水富栄養化防止策として、製油所廃水処理設備からの硫化水素とアンモニア混合ガス含有排水から、硫化水素とアンモニアを分離し、アンモニアガスを回収する方法が開発された。
その一例として、特許文献１に示すAmmonia Recovery Unit(以下ARUプロセスと言う)と呼ばれる技術がある。このARUプロセスの処理フローを図７に示す。
図７において、ARUプロセスは、既存装置にアンモニアスクラバー０6と硫化水素ストリッパー０８を新設するもので、この方法では、既存の廃水処理ストリッパー０１からの塔頂ガスをアンモニアスクラバー０6の最下段にいれる。この既存の塔頂ガスは、約30mol%のアンモニアと約70mol%の硫化水素を含有する。アンモニアスクラバー０6の塔頂から出てくるガスにボイラー供給水を加え冷却器０７により冷却する。その後、ガスと水の混合物を還流ドラム０９に受ける。この還流ドラムに溜まる液は、アンモニア水であり還流ポンプにてリフラックスとしてアンモニアスクラバー０６の塔頂に戻す。還流ドラム０９から出るガスをアンモニアガスＮＨ３として回収する。この方法によれば、製油所のFoul Waterからある程度の純度のアンモニアガス回収は可能である。
ところが、アンモニアスクラバー０６のトレイ上では中和反応が起こり、スクラバーフィード中の硫化水素は水硫化アンモニウム(NH4SH)として固定される。
アンモニアスクラバー０６の塔底から排出される排水には多量の水硫化アンモニウムが残存するので、この水硫化アンモニウム含有水を再度、新設ストリッパーにより硫化水素と残存アンモニアを廃水から分離する。
図中０５、０１１は塔内蒸留温度を所定温度に加熱する熱交換器かまたはストリッピングスチームである。
特開2000-233116号公報 Conventionally, in the oil refining industry, nitrogen and sulfur in crude oil are hydrogenated and removed during the refining process in order to satisfy product standards such as gasoline and kerosene gas oil regardless of domestic and overseas.
During the hydrotreating process, nitrogen and sulfur are converted into ammonia (NH 3 ) and hydrogen sulfide (H 2 S), respectively. Normally, this gas containing hydrogen sulfide and by-product ammonia is washed with water, absorbed in washing water (liquid phase), treated as waste water in a waste water treatment facility (Foul Water Treating Unit), and mixed gas of NH 3 and H 2 S from waste water Is separated by distillation.
FIG. 6 shows a flow of a general known wastewater treatment facility in a refinery.
In this wastewater treatment facility, a mixed gas of NH 3 and H 2 S coming from the top of the existing wastewater stripper 01 of the distillation tower is sent to a sulfur recovery unit (SRU) 04 via a cooler 02 and a reflux drum 03, where The hydrogen sulfide is commercialized as solid sulfur.
Ammonia burns in a heating furnace (not shown) of the sulfur recovery unit (SRU) 04 and is released into the atmosphere as nitrogen oxide gas, while remaining in the treated waste water and released into the sea.
For this reason, the problem of the air pollution by this nitrogen oxide gas discharge | release, the seawater eutrophication by the seawater discharge | release of ammonia, and the problem of generation | occurrence | production of red tide were raised.
Therefore, as a measure to prevent air pollution and seawater eutrophication, a method was developed to separate hydrogen sulfide and ammonia from the hydrogen sulfide and ammonia mixed gas-containing wastewater from the refinery wastewater treatment facility and recover ammonia gas.
As an example, there is a technique called Ammonia Recovery Unit (hereinafter referred to as ARU process) shown in Patent Document 1. The processing flow of this ARU process is shown in FIG.
In FIG. 7, the ARU process is to newly install an ammonia scrubber 06 and a hydrogen sulfide stripper 08 in an existing apparatus. In this method, the tower top gas from the existing wastewater treatment stripper 01 is placed in the lowermost stage of the ammonia scrubber 06. This existing overhead gas contains about 30 mol% ammonia and about 70 mol% hydrogen sulfide. Boiler feed water is added to the gas coming out from the top of the ammonia scrubber 06 and cooled by the cooler 07. Thereafter, the mixture of gas and water is received by the reflux drum 09. The liquid accumulated in the reflux drum is ammonia water and is returned to the top of the ammonia scrubber 06 as a reflux by a reflux pump. The gas exiting from the reflux drum 09 is recovered as ammonia gas NH 3 . According to this method, it is possible to recover ammonia gas with a certain degree of purity from foul water in a refinery.
However, a neutralization reaction occurs on the tray of the ammonia scrubber 06, and hydrogen sulfide in the scrubber feed is fixed as ammonium hydrosulfide (NH 4 SH).
Since a large amount of ammonium hydrosulfide remains in the waste water discharged from the bottom of the ammonia scrubber 06, the hydrogen sulfide-containing water is again separated from the waste water by a new stripper.
In the figure, 05 and 011 are heat exchangers or stripping steam for heating the distillation temperature in the column to a predetermined temperature.
JP 2000-233116 A

このARUプロセスでは、新設されるアンモニアスクラバー０６の塔頂ガス及び硫化水素ストリッパー０８の塔頂に、別途ボイラー給水(純水)０１０を注入するため、全排水量が増大する。このことは、排水の総量規制上好ましくない。また、残存アンモニアを廃水から繰り返し分離するためにエネルギーを消費する。即ち、一度既存廃水ストリッパー01にて蒸留分離した硫化水素とアンモニアを新設アンモニアスクラバー０６でNH4SHとして固定し、再度新設硫化水素ストリッパー０８にて蒸留分離するという非効率さがある。 In this ARU process, since the boiler feed water (pure water) 010 is separately injected into the tower top gas of the ammonia scrubber 06 and the tower top of the hydrogen sulfide stripper 08, the total amount of waste water increases. This is not preferable in terms of regulation of the total amount of waste water. In addition, energy is consumed to repeatedly separate residual ammonia from wastewater. That is, there is an inefficiency that hydrogen sulfide and ammonia once separated by distillation in the existing waste water stripper 01 are fixed as NH 4 SH by the new ammonia scrubber 06 and again distilled by the new hydrogen sulfide stripper 08.

上記問題点を解決するために本発明を、開発したものであり、その特徴とするところは次の通りである。
即ち本発明の特徴とする技術手段は、石油精製装置の重質油脱硫装置から排出した脱硫処理廃水を複数段の蒸留塔に順次供給し、上流の蒸留塔を硫化水素ストリッパーとし、下流の蒸留塔をアンモニア回収ストリッパーとし、硫化水素ストリッパーの塔頂から硫化水素を、アンモニア回収ストリッパーの塔頂から塔頂還流ドラムを介して副生アンモニアを回収する脱硫処理廃水からのアンモニアと硫化水素の分離回収方法において、前記硫化水素ストリッパー塔底の熱交換器(リボイラー)への入熱を変動させることにより前記アンモニア回収ストリッパー塔頂還流ドラムから回収するアンモニア中のH 2 S濃度を所定値に維持する一方、前記アンモニア回収ストリッパーの塔底水を冷却器により所定温度にして、且つ、前記硫化水素ストリッパーの塔頂から出る硫化水素ガス中のアンモニア流量を所定値に維持するため、前記アンモニア回収ストリッパー塔底水の前記冷却器出口の処理後排水の一部を循環ストリームとして分離し、前記循環ストリーム流量を制御し前記硫化水素ストリッパーの塔頂に循環すると共に、前記アンモニア回収ストリッパーの塔頂還流にボイラー給水(純水)を注入することなくアンモニア回収ストリッパーの塔頂還流ドラムから分離した液体をアンモニア回収ストリッパーの塔頂へ循環しながら、前記塔頂還流ドラムからアンモニアガスを回収することを特徴とする脱硫処理廃水からのアンモニアと硫化水素の分離回収方法である。 The present invention has been developed to solve the above problems, and the features thereof are as follows.
That is, the technical means featured in the present invention is to sequentially supply desulfurization treatment waste water discharged from the heavy oil desulfurization unit of the oil refinery unit to a multi-stage distillation column, the upstream distillation column as a hydrogen sulfide stripper, and a downstream distillation unit. Separation and recovery of ammonia and hydrogen sulfide from desulfurization wastewater that recovers hydrogen sulfide from the top of the hydrogen sulfide stripper and by-product ammonia from the top of the ammonia recovery stripper via the tower reflux drum. In the method, the H 2 S concentration in the ammonia recovered from the ammonia recovery stripper tower top reflux drum is maintained at a predetermined value by changing the heat input to the heat exchanger (reboiler) at the bottom of the hydrogen sulfide stripper tower. The bottom water of the ammonia recovery stripper is brought to a predetermined temperature by a cooler, and the hydrogen sulfide stripper In order to maintain the ammonia flow rate in the hydrogen sulfide gas exiting from the top of the column at a predetermined value, a part of the waste water after treatment at the cooler outlet of the ammonia recovery stripper tower bottom water is separated as a circulation stream, and the circulation stream flow rate The liquid separated from the top reflux drum of the ammonia recovery stripper without injecting boiler feed water (pure water) into the top reflux of the ammonia recovery stripper. A method for separating and recovering ammonia and hydrogen sulfide from desulfurization wastewater, wherein ammonia gas is recovered from the top reflux drum while circulating to the top of a stripper.

本発明は、上記構成により、前記脱硫処理廃水からの、アンモニアと硫化水素の分離回収は、ボイラー給水(純水)を注入することなく、前記複数の蒸留塔を用いて、最上流の蒸留塔の塔頂からは硫化水素ガスを高純度で回収し、最下流の蒸留塔の塔頂からアンモニアガスを高純度で且つエネルギー使用の観点から高効率で前記アンモニアガスを回収してその有効利用を有利に実現し、同時に、製油所から大気や海水に排出されるアンモニアを皆無に近く激減させるものである。
また、前記従来技術(ARUプロセス)と違い、本発明においては系内へボイラー給水(純水)を注入しないため、製油所から海に放出される排水量を増大せずにすむものである。 According to the present invention, according to the above configuration, ammonia and hydrogen sulfide can be separated and recovered from the desulfurization treatment wastewater by using the plurality of distillation towers without injecting boiler feed water (pure water). The hydrogen sulfide gas is recovered from the top of the column with high purity, and the ammonia gas is recovered from the top of the most downstream distillation column with high purity and high efficiency from the viewpoint of energy use. It is realized in an advantageous manner, and at the same time, the ammonia discharged from the refinery into the atmosphere and seawater is almost completely reduced.
Further, unlike the above-described prior art (ARU process), in the present invention, boiler feed water (pure water) is not injected into the system, so that it is not necessary to increase the amount of drainage discharged from the refinery to the sea.

本発明において、硫化水素ストリッパー１の設備構成は、多段のダウンカマー付きトレイを内設し底部に貯液部を有する蒸留塔本体１と、蒸留用熱交換器７と、硫黄回収装置１０と、アンモニア回収ストリッパーからのリサイクル流量制御器（Controller−2）１２とからなる。
本発明において、アンモニア回収ストリッパー２の設備構成は、水を頂部から受ける多段のダウンカマー付きトレイを内設し底部に貯液部を有する蒸留塔本体１と、冷却器３，９と、環流ドラム６と、リボイラー熱交換器かまたはストリッピングスチーム８と熱交換器の入熱制御器１３とアンモニアガス回収装置１１とからなる。

又、硫化水素ストリッパー１の好ましい塔圧は、0.7〜1.0MPaの範囲に制御し、アンモニア回収ストリッパー２の好ましい塔圧は、常圧に調整する。
その他、本発明を実施するための好ましい設備構成とプロセスは、次の実施例と共に詳細に説明する。 In the present invention, the equipment configuration of the hydrogen sulfide stripper 1 includes a distillation column main body 1 having a multistage tray with a downcomer and a liquid storage part at the bottom, a distillation heat exchanger 7, a sulfur recovery device 10, It consists of a recycle flow rate controller (Controller-2) 12 from the ammonia recovery stripper.
In the present invention, the equipment configuration of the ammonia recovery stripper 2 includes a distillation column main body 1 having a multi-stage downcomer tray for receiving water from the top and having a liquid storage part at the bottom, coolers 3 and 9, and a recirculation drum. 6, a reboiler heat exchanger or stripping steam 8, a heat input controller 13 of the heat exchanger, and an ammonia gas recovery device 11.

The preferable tower pressure of the hydrogen sulfide stripper 1 is controlled in the range of 0.7 to 1.0 MPa, and the preferable tower pressure of the ammonia recovery stripper 2 is adjusted to normal pressure.
In addition, the preferable equipment configuration and process for carrying out the present invention will be described in detail together with the following examples.

図１に示すプロセスは、製油所における石油精製装置の重質油脱硫装置から排出した脱硫処理廃水(Foul Water)４を少なくとも２段以上の蒸留塔式で処理するプロセスであり、設備構成として30段の場合、第１段から15段のダウンカマー付トレイを内設した蒸留塔を硫化水素ストリッパー1とし、第16段から30段のダウンカマー付トレイを内設した蒸留塔をアンモニア回収ストリッパー2とする。
前記脱硫処理廃水４は、硫化水素ストリッパー1の中段トレイに連続的に送られる。硫化水素ストリッパー1の塔底液は、配管14を介してアンモニア回収ストリッパー2の中段にトレイに連続して圧送される。アンモニア回収ストリッパー2は、塔底水(処理後排水)の一部を海に放流する前に、配管5を介して連続的に硫化水素ストリッパー１の塔頂にリサイクルバックする。各硫化水素ストリッパー1の塔圧は0.7〜1.0MPaの範囲に制御し、各アンモニア回収ストリッパーの塔圧は常圧に調整する。
The process shown in FIG. 1 is a process in which desulfurization treatment waste water (Foul Water) 4 discharged from a heavy oil desulfurization unit of an oil refinery at a refinery is treated with at least two stages of distillation towers. In the case of the stage, the distillation column with the first to 15th tray with downcomer is the hydrogen sulfide stripper 1, and the distillation column with the 16th to 30th tray with downcomer is the ammonia recovery stripper 2. And
The desulfurization waste water 4 is continuously sent to the middle tray of the hydrogen sulfide stripper 1. The column bottom liquid of the hydrogen sulfide stripper 1 is continuously pumped to the tray through the pipe 14 to the middle stage of the ammonia recovery stripper 2. The ammonia recovery stripper 2 is continuously recycled back to the top of the hydrogen sulfide stripper 1 via the pipe 5 before discharging a part of the tower bottom water (the treated waste water) to the sea. The tower pressure of each hydrogen sulfide stripper 1 is controlled in the range of 0.7 to 1.0 MPa, and the tower pressure of each ammonia recovery stripper is adjusted to normal pressure.

この処理フローにより、硫化水素ガスは、各硫化水素ストリッパー1の塔頂から硫黄回収装置１０に回収して固形硫黄として製品化処理する。
アンモニア回収ストリッパー1の塔頂からのアンモニアガスは、冷却器３を介して還流ドラム６に収容しここから高純度のアンモニアガスをアンモニア回収装置１１で回収する。
そして、アンモニア回収ストリッパー2の塔底液(処理後排水)を海に放流する前に配管５を介して硫化水素ストリッパー１の塔頂に返すこと(循環水)により、ボイラー給水を注入する必要が無く、その分製油所の総排水量を低減することになる。
プロセスシミュレーターであるPRO-II/PROVISIONを用いた検証結果では、前記配管５から硫化水素ストリッパー1への循環水の循環量を2ton/hr程度に調節すればアンモニアガスの硫化水素ガスへのスリップを0.1mol/hr程度と最低限にとどめることが可能である。また、硫化水素ストリッパー1の熱交換器(リボイラー)７への入熱を1.0MKcal/hr程度に調節すればアンモニア回収ストリッパー2の塔頂ドラム７から回収されるアンモニアガスの純度は98mol%(2mol%硫化水素ガス)程度まで上げることが可能となる。
<本発明の実施例で使用したシミュレーションツール>
本発明が対象とする石油精製装置の重質油脱硫装置から排出した脱硫処理廃水
系は、水中でアンモニアと硫化水素の中和反応が進行しているので、他のハイドロカーボン系のように沸点の差で精留分離できない。 Through this processing flow, the hydrogen sulfide gas is recovered from the top of each hydrogen sulfide stripper 1 to the sulfur recovery device 10 and commercialized as solid sulfur.
Ammonia gas from the top of the ammonia recovery stripper 1 is accommodated in the reflux drum 6 through the cooler 3, and high purity ammonia gas is recovered from the ammonia drum 11 from there.
And it is necessary to inject boiler feed water by returning to the tower top of the hydrogen sulfide stripper 1 through the pipe 5 (circulated water) before discharging the tower bottom liquid (treated waste water) of the ammonia recovery stripper 2 to the sea. Therefore, the total amount of wastewater from the refinery will be reduced accordingly.
According to the verification results using PRO-II / PROVISION, which is a process simulator, if the circulation rate of the circulating water from the pipe 5 to the hydrogen sulfide stripper 1 is adjusted to about 2 ton / hr, the slip of ammonia gas to hydrogen sulfide gas will occur. It can be kept to a minimum of about 0.1 mol / hr. If the heat input to the heat exchanger (reboiler) 7 of the hydrogen sulfide stripper 1 is adjusted to about 1.0 MKcal / hr, the purity of the ammonia gas recovered from the top drum 7 of the ammonia recovery stripper 2 is 98 mol% (2 mol % Hydrogen sulfide gas).
<Simulation tool used in the embodiment of the present invention>
The desulfurization wastewater system discharged from the heavy oil desulfurization unit of the oil refinery targeted by the present invention has a boiling point like other hydrocarbon systems because the neutralization reaction of ammonia and hydrogen sulfide proceeds in water. The rectification separation is not possible due to the difference of.

そこで、リサイクル系を有する複雑なプロセスのシミュレーションをThermodinamics Method(熱力学的物性推算法)としてGPSWATER(後述)を用いて安定して容易に計算できるツールであるインベンシスシムサイ社の「PRO-II/PROVISION(version 5.55)」を選択した。
同プログラムは、 1990年代に開発されたツールで計算精度、収束性ともに非常によいとされており、 更に同プログラムは蒸留塔をはじめ熱交換器等の他の Unit Operationとネットワークループが組める。また、同プログラムは他のシミュレーションツールと比較しても収束性に遜色がない。 Therefore, the simulation of complex processes with recycle system is a Thermodynamic Properties Method (thermodynamic property estimation method) using GPSWATER (described later). / PROVISION (version 5.55) ”was selected.
This program is a tool developed in the 1990s and is said to have very good calculation accuracy and convergence. Furthermore, this program can be combined with other unit operations such as distillation towers and heat exchangers in a network loop. In addition, the program is comparable to other simulation tools in terms of convergence.

＜熱力学物性値推算法＞
脱硫処理廃水系は、水中にアンモニア、硫化水素、二酸化炭素の電解質を含む系で、気液平衡の計算に必要な熱力学物性値の公知の推算ツールには、「SOUR Method」と「GPSWATER Method」の二つしかないが、本発明の研究では、計算精度の高いとされている「GPSWATER Method」を熱力学物性推算法として採用し、リサイクルループを有するプロセスのシミュレーションをPRO-II/PROVISION(Version5.55)を用いて行った。 <Thermodynamic property value estimation method>
The desulfurization wastewater system contains ammonia, hydrogen sulfide, and carbon dioxide electrolytes in water. Known estimation tools for thermodynamic properties required to calculate vapor-liquid equilibrium include SOUR Method and GPSWATER Method. However, in the research of the present invention, the GPSWATER Method, which is considered to have high calculation accuracy, is adopted as a thermodynamic property estimation method, and simulation of a process having a recycling loop is performed using PRO-II / PROVISION ( Version 5.55).

＜本発明で使用した製油所構内の廃水４（Foul Water）の性状について＞
本実施例では、廃水の組成としては含有窒素分がもっとも多いケースを使用した。イラニアン系の原油を約75%含む原油を処理しているときに実測した重油水素化分解装置(Hydro Cracker)の洗浄後の廃水の性状を使用した。その成分は、H2S濃度：2.08 mol %、NH3濃度：1.22 mol%、その他の成分は、H2Oだけとした。また、フィード流量は20ton/hrとした。 <Properties of waste water 4 (Foul Water) in refinery used in the present invention>
In this example, the case of containing the most nitrogen content was used as the wastewater composition. The properties of waste water after washing in a heavy oil hydrocracking unit (Hydro Cracker) measured when processing crude oil containing about 75% of Iranian crude oil were used. The components were H 2 S concentration: 2.08 mol%, NH 3 concentration: 1.22 mol%, and the other components were H 2 O only. The feed flow rate was 20 ton / hr.

<副生アンモニア回収新プロセスのシミュレーションアルゴリズム>
本実施例でのシミュレーションアルゴリズムを図２に示す。
まず、蒸留塔内のトレイの理論設置段数決定の試算について述べる。
理論段数を決定するためには、種々の因子を考慮し試行錯誤して最終決定しなければならない。
本実施例における、硫化水素ストリッパー内のトレイの理論段の決定因子は、塔頂ガス(H2S)へのアンモニアのスリップを0.1mol/hr程度(Minimum)となるアンモニア回収ストリッパー2塔底からのリサイクル流量が適正な範囲にあるかどうかを主眼とした。
トレイ理論段の初期値として『FRI(Fractionation Institute)が一般的に公表している図6にしめす廃水処理蒸留塔の段効率が40％であること及び製油所で実際に稼動している廃水処理蒸留塔の実トレイ段数が30段程度であること』から15段と設定してシミュレーションを実施した。
その結果該アンモニア回収ストリッパー2塔底から硫化水素ストリッパー1への循環水5の流量が2Ton/hrと計算された。この循環水流量は、フィード流量の20Ton/hrを超えるものでもなくまた、通常のコントロールバルブで流量制御可能な流量である。
従って、硫化水素ストリッパーの理論トレイ段数の15段は、ほぼ適正値であることと判断される。従って、硫化水素ストリッパー内トレイの理論段は15段と決定した。 <Simulation algorithm for new by-product ammonia recovery process>
A simulation algorithm in this embodiment is shown in FIG.
First, a trial calculation for determining the theoretical number of trays in the distillation column will be described.
In order to determine the number of theoretical plates, the final determination must be made by trial and error in consideration of various factors.
In this example, the determinant of the theoretical stage of the tray in the hydrogen sulfide stripper is that the ammonia slip to the tower top gas (H 2 S) slips from the bottom of the ammonia recovery stripper 2 which becomes about 0.1 mol / hr (Minimum). The main focus was on whether the recycle flow rate was within an appropriate range.
As the initial value of the tray theoretical plate, the wastewater treatment distillation tower shown in Fig. 6 which is generally announced by FRI (Fractionation Institute) has a stage efficiency of 40% and wastewater treatment actually operated at the refinery. The simulation was carried out by setting the number of trays in the distillation tower to about 15 ”.
As a result, the flow rate of the circulating water 5 from the bottoms of the two ammonia recovery strippers to the hydrogen sulfide stripper 1 was calculated to be 2 Ton / hr. This circulating water flow rate does not exceed the feed flow rate of 20 Ton / hr, and is a flow rate that can be controlled with a normal control valve.
Accordingly, it is determined that 15 theoretical tray stages of the hydrogen sulfide stripper are almost appropriate values. Therefore, the theoretical stage of the hydrogen sulfide stripper tray was determined to be 15 stages.

一方、本実施例におけるアンモニア回収ストリッパー内トレイの理論段の決定因子はストリッピングスチーム8の流量が『塔底から海に放流される処理後排水中のNH3濃度10ppm以下であることと冷却器９出口温度が40℃という制限』を守れることに主眼をおいた。理論段数の初期値として硫化水素ストリッパーと同様に15段とした。
そのシュミレーション結果、ストリッピングスチーム8の流量が14.4Ton/hrであり、この時のアンモニア回収ストリッパー2の塔底液のNH3濃度は6×10-3 ppm , 冷却器９出口温度は40℃に収束した。このように該ストリッパー2塔底液中のNH3濃度は、規制値の10ppmより低い。従ってアンモニア回収ストリッパー内トレイの理論段は15段で充分であると判断した。 On the other hand, the determinant of the theoretical stage of the tray in the ammonia recovery stripper in this example is that the flow rate of the stripping steam 8 is “the NH 3 concentration in the treated wastewater discharged from the tower bottom to the sea is 10 ppm or less and the cooler The focus was on keeping the 9 outlet temperature limit of 40 ° C. The initial value of the number of theoretical plates was 15 as in the case of the hydrogen sulfide stripper.
As a result of the simulation, the flow rate of the stripping steam 8 is 14.4 Ton / hr, the NH 3 concentration of the bottom liquid of the ammonia recovery stripper 2 at this time is 6 × 10 −3 ppm, and the outlet temperature of the cooler 9 is 40 ° C. Converged. Thus, the NH 3 concentration in the stripper 2 bottom liquid is lower than the regulated value of 10 ppm. Therefore, it was judged that 15 theoretical plates for the ammonia recovery stripper tray were sufficient.

図２に示す様に本発明で行ったシュミューレーションのロジックは、以下の通りである。
(1). 硫化水素ストリッパーでのスペックとバリアブル（変数）
同塔底水中のH2S の流量が1.1mol/h(Spec1)となるように同塔のリボイラー７の入熱をバリアブル(変数)として変動させた。蒸留塔外部に制御器（Controller-1）１３を設置して上記のSpec 1を変数として変動させ、各アンモニア回収ストリッパー塔頂から回収するアンモニア中のH2S濃度を2mol%に制御した。
この2mol%とした理由は図５で詳述する。
(2). アンモニア回収ストリッパー2でのスペックとバリアブル(変数)
アンモニア回収ストリッパー2のコンデンサー3の出口温度が40℃となるようにコンデンサー3からの出熱量を変動させた。アンモニア回収ストリッパー2塔底出口に冷却器９を設置し同冷却器９出口温度を40℃まで冷却した。
(3). 循環水ストリームの作成
その後、スピリッターを設置し海に放流するストリーム15と硫化水素ストリッパー1塔頂に循環するストリーム5に分けた。
(4). 循環ストリームの流量コントロール
この循環ストリーム流量を変数として各硫化水素ストリッパー塔頂から出る硫化水素ガス中のアンモニア流量が0.1mol/hとなるように即ち、硫化水素ガス中へのアンモニアのスリップが最小となるように制御器（Controller-2）１２にて制御した。 As shown in FIG. 2, the simulation logic performed in the present invention is as follows.
(1). Specification and variable (variable) in hydrogen sulfide stripper
The heat input of the reboiler 7 of the tower was varied as a variable (variable) so that the flow rate of H 2 S in the tower bottom water was 1.1 mol / h (Spec1). A controller (Controller-1) 13 was installed outside the distillation column to vary the above Spec 1 as a variable, and the H 2 S concentration in the ammonia recovered from the top of each ammonia recovery stripper column was controlled to 2 mol%.
The reason for the 2 mol% will be described in detail with reference to FIG.
(2). Specification and variable (variable) in ammonia recovery stripper 2
The amount of heat output from the condenser 3 was varied so that the outlet temperature of the condenser 3 of the ammonia recovery stripper 2 was 40 ° C. A cooler 9 was installed at the bottom outlet of the ammonia recovery stripper 2 tower, and the outlet temperature of the cooler 9 was cooled to 40 ° C.
(3). Creation of a circulating water stream After that, a splitter was installed to separate the stream 15 into the sea and the stream 5 circulating into the top of one hydrogen sulfide stripper.
(4). Flow rate control of the circulating stream With this circulating stream flow rate as a variable, the ammonia flow rate in the hydrogen sulfide gas exiting from the top of each hydrogen sulfide stripper tower is 0.1 mol / h, that is, the ammonia flow into the hydrogen sulfide gas. Control was performed by a controller (Controller-2) 12 so as to minimize the slip.

<シミュレーション結果と考察>
(1). 硫化水素ストリッパーの塔圧の決定
「加熱する熱源温度と加熱される低温側との温度差(ΔT)」に対する
「単位面積あたりの伝熱量」との相関を図３に示した。
図３から読み取れるようにグラフ中の点B―C間では、温度差(ΔT)の増加と伴に単位面積当たりの伝熱量は点A−B間と比較すると急激に大きくなっている。このB−Cの域を核沸騰領域と呼び、蒸発を伴う熱交換器７(リボイラー等)はこの温度差領域で設計する。
さて、製油所内で一般的に使用される蒸留塔の熱交換器７の熱源スチームは圧力が1.2MPaGのスチーム(以下12K STMと呼ぶ)であり、12K STMの飽和温度は191℃である。
硫化水素ストリッパーの塔底温度はその塔圧における飽和温度であるので、硫化水素ストリッパー塔圧は、12K STMを熱源に使用する限り必然的に決まる。即ち、図３の点Cを越えない点B−C間の核沸騰領域に塔底温度が入る塔圧に限定される。核沸騰領域に入る温度差ΔT(加熱媒体の温度−加熱される側の温度)を図３から読み取ると8℃〜20℃である。
本実施例では、硫化水素ストリッパーの圧力を0.4MPaG、0.8MPaG、1.1MPaGと変化させてシミュレーションを行い、予備検討を行った。そのシミュレーション結果から「硫化水素ストリッパーの塔圧」と「12K STMと塔底温度との差」との相関関係を図４に示した。図４から熱交換器７が核沸騰領域に入る塔圧は、0.7MPaG〜0.95MPaGの範囲にあることが分かる。以上のことより本発明では、各硫化水素ストリッパーの塔圧を0.8MPaGに設定した。
(2). 硫化水素ストリッパーの熱原と、アンモニア回収ストリッパー塔頂からのアンモニアガス中の硫化水素濃度。
硫化水素ストリッパーの熱源を増加すれば、アンモニア回収ストリッパーから回収されるアンモニア中の硫化水素濃度は下がる。図５は、硫化水素ストリッパー熱源量とアンモニア回収ストリッパーの塔頂から回収されるアンモニア中の硫化水素濃度との相関関係を示した図である。
図5から回収アンモニア中の硫化水素濃度を1mol%以下にするためには、ほぼ無限大の熱量が必要ということが読み取れる。このグラフより、必要熱源量と上記硫化水素濃度との間で最適な点は、H2S%が2mol%程度と考えられる。
従って、本発明での全てのシミュレーションはこのH2S%を2mol%に固定して行った。
(3). アンモニアと硫化水素濃度分離に必要な熱量(硫化水素ストリッパー熱原)の妥当性評価
現在の実運転データより、廃水処理の蒸留塔に使用されているエネルギーは、3.3Mkcal/hr低度である。一方、NH3/H2S分離に必要なエネルギーは、98mol%NH3(2%molH2S)を得るためには図５より1.0Mkcal/hr程度である。
上記２つの必要エネルギーを比較すると98mol%NH3ガスを廃水から分離し回収するために必要なエネルギーは妥当な量である。 <Simulation results and discussion>
(1). FIG. 3 shows the correlation between the heat transfer amount per unit area with respect to the determination of the tower pressure of the hydrogen sulfide stripper and the “temperature difference (ΔT) between the heat source temperature to be heated and the low temperature side to be heated”.
As can be seen from FIG. 3, between points B and C in the graph, the amount of heat transfer per unit area rapidly increases as the temperature difference (ΔT) increases compared to between points A and B. This B-C region is called the nucleate boiling region, and the heat exchanger 7 (reboiler etc.) accompanied by evaporation is designed in this temperature difference region.
Now, the heat source steam of the heat exchanger 7 of the distillation tower generally used in the refinery is steam having a pressure of 1.2 MPaG (hereinafter referred to as 12K STM), and the saturation temperature of 12K STM is 191 ° C.
Since the bottom temperature of the hydrogen sulfide stripper is the saturation temperature at that tower pressure, the hydrogen sulfide stripper tower pressure is inevitably determined as long as 12K STM is used as the heat source. That is, it is limited to the tower pressure at which the tower bottom temperature enters the nucleate boiling region between points B and C not exceeding the point C in FIG. When the temperature difference ΔT (temperature of the heating medium−temperature on the heated side) entering the nucleate boiling region is read from FIG. 3, it is 8 ° C. to 20 ° C.
In this example, a simulation was performed by changing the pressure of the hydrogen sulfide stripper to 0.4 MPaG, 0.8 MPaG, and 1.1 MPaG, and a preliminary study was performed. From the simulation results, the correlation between “the tower pressure of the hydrogen sulfide stripper” and “the difference between 12K STM and the tower bottom temperature” is shown in FIG. It can be seen from FIG. 4 that the tower pressure at which the heat exchanger 7 enters the nucleate boiling region is in the range of 0.7 MPaG to 0.95 MPaG. From the above, in the present invention, the tower pressure of each hydrogen sulfide stripper was set to 0.8 MPaG.
(2). Hydrogen sulfide concentration in the hydrogen gas stripper heat source and ammonia recovery stripper tower top.
If the heat source of the hydrogen sulfide stripper is increased, the concentration of hydrogen sulfide in the ammonia recovered from the ammonia recovery stripper is lowered. FIG. 5 is a diagram showing the correlation between the amount of heat source of the hydrogen sulfide stripper and the concentration of hydrogen sulfide in the ammonia recovered from the top of the ammonia recovery stripper.
It can be seen from FIG. 5 that an almost infinite amount of heat is required to reduce the hydrogen sulfide concentration in the recovered ammonia to 1 mol% or less. From this graph, the optimum point between the required heat source amount and the hydrogen sulfide concentration is considered that H 2 S% is about 2 mol%.
Therefore, all simulations in the present invention were performed with this H 2 S% fixed at 2 mol%.
(3). Appropriate evaluation of calorific value (hydrogen sulfide stripper heat field) required for ammonia and hydrogen sulfide concentration separation Based on the current actual operation data, the energy used in the distillation column for wastewater treatment is 3.3 Mkcal / hr lower. Degree. On the other hand, the energy required for NH 3 / H 2 S separation is about 1.0 Mkcal / hr from FIG. 5 in order to obtain 98 mol% NH 3 (2% mol H 2 S).
Comparing the above two required energies, the amount of energy required to separate and recover 98 mol% NH 3 gas from wastewater is a reasonable amount.

以上説明したように本発明は、副生アンモニアを燃焼し多量の酸化窒素を大気放出することなく、また、海水にアンモニアを放出することなく、副生アンモニアを98％以上の高純度で回収し資源化することができる。アンモニアガスの価格は、現在70,000円/tonと高価なものであり、本発明による経済効果は非常に大きく、また、環境汚染の一因となっている酸化窒素化合物や海に放出され海水の富栄養化による赤潮の一因となっているアンモニアの放出を最小に低減でき環境保全に大きく貢献するものである。特に、本発明によればアンモニアの精製過程でボイラー給水(純水)を注入する必要がなく製油所から排出される総排水量を増加させずに済む。また、本発明は、副生アンモニア精製に必要なエネルギーは硫化水素とアンモニアの分離のために1.0MKcal/hr程度必要であるがこれは、現在稼動中の廃水処理設備に使用されているエネルギーの3分の1程度であり極めて経済的であるなどの幾多の効果を呈しこの種産業上の利用可能性多大なものがある。 As described above, the present invention recovers by-product ammonia with a high purity of 98% or more without burning by-product ammonia and releasing a large amount of nitric oxide into the atmosphere, and without releasing ammonia into seawater. It can be turned into resources. The price of ammonia gas is currently expensive at 70,000 yen / ton, and the economic effects of the present invention are very large. Also, the richness of seawater released to the sea, such as nitric oxide compounds that contribute to environmental pollution. The release of ammonia, which contributes to the red tide due to nutrition, can be reduced to a minimum and contribute greatly to environmental conservation. In particular, according to the present invention, it is not necessary to inject boiler feed water (pure water) during the purification process of ammonia, and it is not necessary to increase the total amount of drainage discharged from the refinery. In the present invention, the energy required for refining by-product ammonia is about 1.0 MKcal / hr for the separation of hydrogen sulfide and ammonia, which is the energy used for the wastewater treatment facility currently in operation. There are a number of effects such as being about one third and extremely economical, and there is a great deal of industrial applicability.

製油所における石油精製装置の重質油脱硫装置から排出した脱硫処理廃水４を二基の蒸留塔で処理するプロセスを示す説明図。Explanatory drawing which shows the process which processes the desulfurization process waste water 4 discharged | emitted from the heavy oil desulfurization apparatus of the oil refinery apparatus in a refinery with two distillation towers. 本発明のシミュレーションアルゴリズム（シュミューレーションのロジック）を示す説明図。Explanatory drawing which shows the simulation algorithm (logic of simulation) of this invention. 「加熱する熱源温度と加熱される低温側との温度差(ΔT)」に対する「単位面積あたりの伝熱量』との相関図。（核沸騰領域に入る温度差ΔT(Hot Side Temp.−Cold Side Temp.)読み取り図）Correlation diagram of “heat transfer amount per unit area” with respect to “temperature difference between heated heat source temperature and heated low temperature side (ΔT)” (temperature difference ΔT (Hot Side Temp.-Cold Side (Temp.) Reading) 『第1蒸留塔の塔圧』と『12K STMと蒸留塔底温度との差』との相関関係図。Correlation diagram between "column pressure of first distillation column" and "difference between 12K STM and distillation column bottom temperature". 硫化水素ストリッパー１熱源量とアンモニア回収ストリッパー２の塔頂から回収されるアンモニア中の硫化水素濃度との相関関係を示した図。The figure which showed the correlation with the hydrogen sulfide concentration in the ammonia collect | recovered from the tower | column top of the hydrogen sulfide stripper 1 heat source amount and the ammonia collection | recovery stripper 2. FIG. 製油所における従来の一般的な廃水処理設備フロー図。The conventional general wastewater treatment facility flow chart in a refinery. 特許文献１に示すAmmonia Recovery Unit(以下ARUプロセスと言う)と呼ばれる処理フロー図。FIG. 6 is a process flow diagram called Ammonia Recovery Unit (hereinafter referred to as ARU process) shown in Patent Document 1;

符号の説明Explanation of symbols

１ 硫化水素ストリッパー
２ アンモニア回収ストリッパー
３、９ 冷却器
４ 脱硫処理廃水
６ 環流ドラム
７、８ 熱交換器
１０ 硫黄回収装置
１１ アンモニア液回収装置
１２、１３ 制御器
DESCRIPTION OF SYMBOLS 1 Hydrogen sulfide stripper 2 Ammonia collection | recovery stripper 3, 9 Cooler 4 Desulfurization waste water 6 Circulation drum 7, 8 Heat exchanger 10 Sulfur recovery apparatus 11 Ammonia liquid recovery apparatus 12, 13 Controller

Claims (1)

石油精製装置の重質油脱硫装置から排出した脱硫処理廃水を複数段の蒸留塔に順次供給し、上流の蒸留塔を硫化水素ストリッパーとし、下流の蒸留塔をアンモニア回収ストリッパーとし、硫化水素ストリッパーの塔頂から硫化水素を、アンモニア回収ストリッパーの塔頂から塔頂還流ドラムを介して副生アンモニアを回収する脱硫処理廃水からのアンモニアと硫化水素の分離回収方法において、前記硫化水素ストリッパー塔底の熱交換器(リボイラー)への入熱を変動させることにより前記アンモニア回収ストリッパー塔頂還流ドラムから回収するアンモニア中のH 2 S濃度を所定値に維持する一方、前記アンモニア回収ストリッパーの塔底水を冷却器により所定温度にして、且つ、前記硫化水素ストリッパーの塔頂から出る硫化水素ガス中のアンモニア流量を所定値に維持するため、前記アンモニア回収ストリッパー塔底水の前記冷却器出口の処理後排水の一部を循環ストリームとして分離し、前記循環ストリーム流量を制御し前記硫化水素ストリッパーの塔頂に循環すると共に、前記アンモニア回収ストリッパーの塔頂還流にボイラー給水(純水)を注入することなくアンモニア回収ストリッパーの塔頂還流ドラムから分離した液体をアンモニア回収ストリッパーの塔頂へ循環しながら、前記塔頂還流ドラムからアンモニアガスを回収することを特徴とする脱硫処理廃水からのアンモニアと硫化水素の分離回収方法。 The desulfurization treatment wastewater discharged from the heavy oil desulfurization unit of the oil refining unit is sequentially supplied to the multistage distillation column, the upstream distillation column is the hydrogen sulfide stripper, the downstream distillation column is the ammonia recovery stripper, and the hydrogen sulfide stripper hydrogen sulfide from the column top, the method of separating and recovering ammonia and hydrogen sulfide from the desulfurization wastewater to recover byproduct ammonia via the top reflux drum from the top of the ammonia recovery stripper, of the hydrogen sulfide stripper bottoms thermal By changing the heat input to the exchanger (reboiler), the H 2 S concentration in the ammonia recovered from the ammonia recovery stripper top reflux drum is maintained at a predetermined value, while the bottom water of the ammonia recovery stripper is cooled. In the hydrogen sulfide gas coming out from the top of the hydrogen sulfide stripper. In order to maintain the monia flow rate at a predetermined value, a part of the treated waste water from the cooler outlet of the ammonia recovery stripper tower bottom water is separated as a circulation stream, and the circulation stream flow rate is controlled to control the top of the hydrogen sulfide stripper. Circulated to the top of the ammonia recovery stripper while circulating the liquid separated from the top reflux drum of the ammonia recovery stripper without injecting boiler feed water (pure water) to the top reflux of the ammonia recovery stripper. A method for separating and recovering ammonia and hydrogen sulfide from desulfurization wastewater, wherein ammonia gas is recovered from a tower top reflux drum .
JP2004367675A 2004-12-20 2004-12-20 A method for separating and recovering ammonia and hydrogen sulfide from desulfurization wastewater. Expired - Fee Related JP4573106B2 (en)

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